Method of producing lower alcohols from glycerol

ABSTRACT

A reactive-separation process converts glycerin into lower alcohols, having boiling points less than 200° C., at high yields. Conversion of natural glycerin to propylene glycol through an acetol intermediate is achieved at temperatures from 150° to 250° C. at pressures from 1 and 25 bar. The preferred applications of the propylene glycol are as an antifreeze, deicing compound, or anti-icing compound. The preferred catalyst for this process in a copper-chromium.

RELATED APPLICATIONS

This application claims benefit of priority to U.S. provisionalapplication Ser. No. 60/731,673 filed Oct. 31, 2005 and is acontinuation-in-part of U.S. application Ser. No. 11/088,603 filed Mar.24, 2005, now U.S Pat. No. 7,663,004 which claims benefit of priority toU.S. provisional patent application Ser. No. 60/556,334 filed Mar. 25,2004 and is a continuation-in-part of U.S. patent application Ser. No.10/420,047 filed Apr. 21, 2003, now abandoned which claims benefit ofpriority to U.S. provisional patent application: Ser. Nos. 60/374,292,filed Apr. 22, 2002 and 60/410,324, filed Sep. 13, 2002, all of whichare incorporated by reference herein.

BACKGROUND

1. Field of the Invention

This invention relates generally to a process for value-added processingof fats and oils to yield glycerol and glycerol derivatives. Moreparticularly, the process converts glycerol to acetol and/or propyleneglycol, which is also known as 1, 2 propanediol. The process may yieldglycerol-based products and glycerol derivatives, such as antifreeze andother products.

2. Description of the Related Art

Existing processes for the hydrogenation of glycerol to form otherproducts are generally characterized by requirements for excessivelyhigh temperatures and pressures. For example, high temperatures maydegrade the reaction products. Working pressures of several hundred barcreate safety concerns and increase the capital costs of implementingthese processes. Most of such processes yield substantial impuritiesthat may necessitate costly purification steps to isolate the desiredreaction products.

In one example, conventional processing of natural glycerol topropanediols uses a catalyst, for example, as reported in U.S. Pat. Nos.5,616,817, 4,642,394, 5,214,219 and U.S. Pat. No. 5,276,181. Thesepatents report the successful hydrogenation of glycerol to formpropanediols. None of the processes shown by these patents provide adirect reaction product mixture that is suitable for use as antifreeze.None provide process conditions and reactions that suitably optimize theresultant reaction product mixture for direct use as antifreeze. Noneaddress the use of unrefined crude natural glycerol feed stock, and noneof these processes are based on reactive distillation. Generally,existing processes

U.S. Pat. No. 5,616,817 issued to Schuster et al. describes thecatalytic hydrogenation of glycerol to produce propylene glycol in highyield, such as a 92% yield, with associated formation of n-propanol andlower alcohols. Conversion of glycerol is substantially complete using amixed catalyst of cobalt, copper, manganese, and molybdenum.Hydrogenation conditions include a pressure of from 100 to 700 bar and atemperature ranging from 180° C. to 270° C. Preferred process conditionsinclude a pressure of from 200 to 325 bar and a temperature of from 200°C. to 250° C. This is because Schuster et al. determined that lowerpressures lead to incomplete reactions, and the higher pressuresincreasingly form short chain alcohols. A crude glycerol feed may beused, such as is obtainable from the transesterification of fats andoils, but needs to be refined by short path distillation to removecontaminants, such as sulfuric acid that is commonly utilized in thetransesterification process. The feed should contain glycerol in highpurity with not more than 20% water by weight.

U.S. Pat. No. 4,642,394 issued to Che et al. describes a process forcatalytic hydrogenation of glycerol using a catalyst that containstungsten and a Group VIII metal. Process conditions include a pressureranging from 100 psi to 15,000 psi and a temperature ranging from 75° C.to 250° C. Preferred process conditions include a temperature rangingfrom 100° C. to 200° C. and a pressure ranging from 200 to 10,000 psi.The reaction uses basic reaction conditions, such as may be provided byan amine or amide solvent, a metal hydroxide, a metal carbonate, or aquaternary ammonium compound. The concentration of solvent may be from 5to 100 ml solvent per gram of glycerol. Carbon monoxide is used tostabilize and activate the catalyst. The working examples show thatprocess yields may be altered by using different catalysts, for example,where the yield of propanediols may be adjusted from 0% to 36% basedupon the reported weight of glycerol reagent.

U.S. Pat. Nos. 5,214,219 issued to Casale, et al. and 5,266,181 issuedto Matsumura, et al. describe the catalytic hydrogenation of glycerolusing a copper/zinc catalyst. Process conditions include a pressureranging from 5 MPa to 20 MPa and a temperature greater than 200° C.Preferred process conditions include a pressure ranging from 10 to 15MPa and a temperature ranging from 220° C. to 280° C. The concentrationof glycerol may range from 20% to 60% by weight in water or alcohol, andthis is preferably from 30% to 40% by weight. The reaction may beadjusted to produce significant amounts of hydrocarbon gas and/or lacticacid, such that gas generation is high when lactic acid formation is lowand lactic acid formation is high when gas generation is low. Thisdifference is a function of the amount of base, i.e., sodium hydroxide,which is added to the solvent. Alcohol reaction products may range from0% to 13% of hydrocarbon products in the reaction mixture by molarpercentages, and propanediols from 27% to 80%. Glycerol conversionefficiency exists within a range from 6% to 100%.

SUMMARY

The presently disclosed instrumentalities advance the art and overcomesthe problems outlined above by producing value-added products inexceptionally high yield and purity from hydrogenation of naturalglycerol feed stocks. In other aspects, the disclosure pertains to themanufacture of products that do not require exceptionally high yield andpurity, such as antifreeze.

In one aspect, a process for converting glycerol to acetol with highselectivity, commences with providing a glycerol-containing materialthat has 50% or less by weight water. This material may be, for example,a byproduct of biodiesel manufacture. The glycerol-containing materialis contacted with a catalyst that is capable of hydrogenating glycerol,in order to form a reaction mixture. Conditions for reaction of thereaction mixture are established to include a temperature within a rangefrom 150° C. to 250° C. and a pressure within a range from 0.1 bar to 25bar. The reaction mixture is reacted under the conditions for reactionto dehydrate the glycerol with resultant formation of acetol as areaction product. The reaction may be performed at temperatures of up to270° C., 280° C. or even 290° C.; however, the use of these increasedtemperature results in thermal degradation of the reaction producttogether with die-reactions, and so is not recommended for applicationswhere high purity of the reaction product is required. It is possible byuse of this process according to one or more of the embodimentsdescribed below to achieve, for example, propylene glycol that is 90% preven 98% pure at a high yield of better than 85% or even 95%.

In various other aspects, the glycerol-containing feedstock preferablycontains from 5% to 15% water by weight. The catalyst may be aheterogenous catalyst that contains at least one element from Groups Ior VIII of the Periodic Table. The catalyst may be a heterogeneouscatalyst including at least one material selected from the groupconsisting of palladium, nickel, rhodium, copper, zinc, chromium andcombinations thereof. The dehydration catalyst may, for example, containfrom 5 wt % to 95 wt % chromium, and may be comprised of compositions ofcopper expressed as CuO and chromium expressed as Cr₂O₃ at 30-80 wt % ofCuO and 20-60 wt % of Cr₂O₃. In one example, the catalyst may beexpressed as Cr₂O₃ at 40-60 wt % of CuO and 40-50 wt % of Cr₂O₃. Thepresence of hydrogen reduces these oxides with their reduced form whichis the active form of the catalyst for hydrogenation of acetol.

A small amount of hydrogen may be added to deter the acetol reactionproduct during formation from scavenging hydrogen from other hydrocarbonmaterials in the reaction mixture. If acetol is the desired finalproduct, the partial pressure of hydrogen may be sufficiently low, suchas about 0.1 bar, to prevent substantial conversion of acetol topropylene glycol.

A greater amount of hydrogen may be added to facilitate conversion ofthe acetol to other products. Where hydrogen is added under theforegoing reaction conditions, the dominant product is suitablypropylene glycol.

It is possible to use a gas flow for stripping the reaction productsfrom the reaction mixture, where such reaction product may includeacetol and propylene glycol. In one embodiment, the glycerol-containingmaterial is in liquid phase and the process entails removing thereaction product(s) during the step of reacting. This may be done byfacilitating selective release of acetol as vapor from the reactionmixture by action of partial pressure through contact with a gas, suchas nitrogen or a noble gas, that is essentially inert to the reactionmixture and the acetol reaction product.

The acetol may be condensed and further reacted to form downstreamproducts, such as by reaction with hydrogen to produce propylene glycolor lactaldehyde. A condenser for this purpose suitably operates at atemperature between 25° C. and 150° C., or more preferably from 25° C.to 60° C. One process for converting acetol to propylene glycol withhigh selectivity entails contacting an acetol-containing feedstock thatcontains less than 50% by weight water with a catalyst that is capableof hydrogenating acetol to form a reaction mixture; and heating thereaction mixture to a temperature between 50° to 250° C. at a pressurebetween 1 and 500 bar to form propylene glycol.

In another embodiment, the gas that strips reaction [products from theinitial reaction mixture may be reactive with the acetol reactionproduct, such as hydrogen gas is reactive with the acetol. Accordingly,the stripper gas may be supplemented with hydrogen for this effect, suchthat a different reaction product is condensed. This different reactionproduct may be propylene glycol. Unused hydrogen may be recycled fromthe condenser back to the reactor vessel.

A more preferred temperature range for facilitating the reaction is from180° C. to 220° C. A more preferred pressure range is from 1 to 20 bar,where low pressures of from 1 to 15 bar and 1 to 5 bar may yieldespecially pure products. The reaction may persist for a duration in aslurry phase with reaction limited by catalyst within a range from 0.1hour to 96 hours, such as from 4 to 46 hours or from 4 to 28 hours. Itis possible to operate the reaction at higher catalyst loadings and evenin a gas phase with much shorter reaction times within the range from0.001 to 8 hours, or more-preferably 0.002 to 1 hour, or even morepreferably from 0.05 to 0.5 hours.

In another embodiment, the reaction does not require a glycerol feed,but may be a polyhydric material, such as a three-carbon or greatersugar or polysaccharide. The process equipment in use on these materialsmay form an alcohol product having a boiling point less than 200° C.

Batch reactor effluent may be used as an antifreeze, anti-icing agent orde-icing agent, for example, as may be obtained from the crude glycerolbyproduct of the C1 to C4 alkyl alcohol alcoholysis of a glyceride. Analternative glycerol source is the crude from hydrolysis of a glyceride.Such materials as this may contain, on a water-free basis, from about0.5% to about 60% glycerol, and from about 20% to about 85% propyleneglycol. Another such composition may contain, on a water-free basis,from about 10% to about 35% glycerol, from about 40% to about 75%propylene glycol, and from about 0.2% to about 10% C1 to C4 alkylalcohol. The compositions may also contain from about 1% to 15% residueby-product from a reaction of glycerol.

In one embodiment, a process for producing antifreeze from a crudeglycerol byproduct of a C₁ to C₄ alkyl alcohol alcoholysis of aglyceride, entails neutralizing the crude glycerol to achieve a pHbetween 5 and 12. This is followed by separating C₁ to C₄ alcohol andwater from the crude glycerol such that the combined concentrations ofwater and C₁ to C₄ alcohols is less than about 5 (wt) %. The separatedcrude glycerol is contacted with a hydrogenation catalyst and hydrogenat a pressure of between about 0.1 and 200 bar and at a temperaturebetween about 100° C. and 280° C. for a period of time sufficient toachieve a conversion of the glycerol of between 60 and 90% on the basisof glycerol in the crude glycerol. The pressure is more preferablywithin a range from 0.1 to 25 bar and is even more preferably from 1 to20 bar. Separation of C₁ to C₄ alcohols and water may be achieved byflash separation at a temperature greater than about 60° C., or bythermal diffusion. The hydrogenation catalyst may contain at least onemetal from the group consisting of palladium, nickel, zinc, copper,platinum, rhodium, chromium, and ruthenium.

A gas phase reaction may be performed for converting glycerol to aproduct at high selectivity to propylene glycol and low selectivity toethylene glycol. The reaction commences with providing a gas phasereaction mixture that is essentially free of liquid and contains:glycerol with a partial pressure of glycerol in a range from 0.01 barsand 0.5 bars of glycerol, and hydrogen with a partial pressure ofhydrogen between 0.01 and 25 bars of hydrogen. The reaction mixture ismaintained at a total pressure between 0.02 and 25 bars and contacts aheterogeneous catalyst at a temperature between 150° C. and 280° C. toform propylene glycol.

In the gas phase reaction, a partial pressure of glycerol is preferablyless than glycerol's dew point partial pressure in the reaction mixture,and greater than one fourth the dew point partial pressure in thereaction mixture. This partial pressure is also preferably greater thanhalf the dew point partial pressure in the reaction mixture. The gasphase reaction mixture contains essentially no liquid and has a partialpressure of glycerol between 0.01 and 0.5 bars of glycerol and a partialpressure of hydrogen between 0.01 and 5 bars of hydrogen; and thereaction may be performed at a temperature between 150° C. and 280° C.to facilitate a reaction by use of the same catalysts described above.The total pressure of reaction may be between 0.02 and 5 bars.

The process may be tuned to produce increased amounts of lactaldehydewith high selectivity. This is done by combining a glycerol-containingfeedstock with less than 50% by weight water with a catalyst that iscapable of dehydrating glycerol to form a reaction mixture; and heatingthe reaction mixture to a temperature between 150° to 200° C. at apressure between 0.01 and 25 bar. A preferred temperature range for thisreaction is from 165° C. to 185° C., while the pressure exists within arange from 0.02 to 2 bars. The lactaldehyde condenser may operate at atemperature between 0° C. to 140° C.

The propylene glycol product may be produced in high purity. especiallyfrom the gas-phase reaction. The propylene glycol reaction product maybe further purified by adding a base to the said propylene glycolproduct to achieve a pH greater than 8.0 and distilling the propyleneglycol from the product having a pH greater than 8.0. The base may beselected from the group comprised of sodium hydroxide, potassiumhydroxide, and calcium oxide.

Although a batch reactor is preferred, other suitable reactor typesinclude slurry batch reactors, trickle bed reactors, and teabagreactors. One reactor for use with highly exothermic reactions comprisedof an outer shell containing U-tubes with an orientation such that theU-end of the U-Tubes is facing upward. The shell has an upper removablehead where catalyst is loaded between shell and tubes from the top byremoving the upper head. An inert packing may be is placed in the lowestportion of the space between the shell and U-Tubes at a depth between 2and 24 inches

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 is a schematic block flow diagram illustrating preferredreactor-separator with a reactor, condenser, and condensate tank, andrecycle of unreacted hydrogen.

FIG. 2 is a schematic of the proposed reaction mechanism for conversionof glycerol to propylene glycol via acetol intermediate.

FIG. 3 is a schematic of the proposed reaction mechanism for conversionof acetol to propylene glycol via lactaldehyde intermediate.

FIG. 4 is a schematic of the proposed two-step alternative embodimentfor converting glycerol to acetol and then converting acetol topropylene glycol, where the process equipment may also be used to makepropylene glycol with no intermediate step;

FIG. 5 is a schematic of laboratory process equipment that may be usedto demonstrate the process equipment of FIG. 4 or 6.

FIG. 6 is a schematic of the proposed two-step alternative embodimentfor converting glycerol to acetol and then converting acetol topropylene glycol where hydrogen is used for the first reactor at a lowerpressure and water is removed from the vapor effluents from the firstreactor to allow purging of the water from the system.

FIG. 7 is a schematic block flow diagram illustrating a packed-bedreactor with an evaporator, reactor, and condenser.

FIG. 8 shows pressure dependence of the glycerol to propylene glycolreaction at temperatures of 220° C., and 240° C.

FIG. 9 shows glycerol to propylene glycol reaction effect of H₂:glycerol mole ratio on catalyst productivity at 220° C.

FIG. 10 shows glycerol to PG reaction: effect of H2: Glycerol mole ratioon catalyst productivity at 220° C.

FIGS. 11A, 11B and 11C show a preferred reactor configuration.

FIG. 12 shows a pilot scale reactor.

FIG. 13 illustrates a packed-bed reactor with optional gas feed toevaporator.

FIG. 14 is a schematic block flow diagram illustrating a preferredpacked-bed reactor system including recycle of product to improvetemperature control and purification of the reactor effluent in aseparator.

FIG. 15 is a schematic block flow diagram illustrating a preferredpacked-bed reactor system including recycle of product to improvetemperature control and purification of the reactor effluent in aseparator

DETAILED DESCRIPTION OF PREFERRED EMBODIMENTS

There will now be shown and described by way of non-limiting example aprocess for producing lower alcohols from glycerol feed stock to provideglycerol-based and/or propylene glycol-based antifreezes. The loweralcohols include, for example, as acetol and propylene glycol. Preferreduses of reaction product mixtures that are derived from the processinclude but are not limited to deicing fluids, anti-icing fluids, andantifreeze applications. These uses of the glycerol-based and/orpropylene glycol-based antifreezes displace the use of toxic andnon-renewable ethylene glycol with non-toxic and renewableglycerol-derived antifreeze. In this regard, use of propylene glycolthat is derived from natural glycerol is a renewable alternative topetroleum-derived propylene glycol. Other downstream uses for propyleneglycol include any substitution or replacement of ethylene glycol orglycerol with propylene glycol.

Equipment for Reactive-Separation Preparation of Antifreeze fromPoly-Alcohols Like Glycerol

One method of preparing antifreeze from glycerol includes reaction at atemperature ranging from 150° to 250° C. and in some embodiments thistemperature is more preferably from 180° C. to 220° C. The reactionoccurs in a reaction vessel. The pressures in the reaction vessel arepreferably from 1 to 25 bars and in some embodiments this pressure ismore-preferably between 5 and 18 bars. The process equipment mayinclude, for example, a reactor at these temperature and pressureconditions connected to a condenser and condensate tank where thecondenser is preferably at a temperature between about 25° C. and 150°C. and in some embodiments this is more preferably between 25° and 60°C.

FIG. 1 provides a block flow diagram of process equipment 100 includinga reactor-separator 102. A polyhydric feed stock 104, for example,containing glycerol, is introduced stepwise or continuously into thereactor separator 102. Hydrogen 106 is added to hydrogen line 108 topromote conversion of glycerol 104 to propylene glycol within thereactor-separator 102. The process temperatures are such that adistillation occurs with the formation or presence of propylene glycol,short chain alcohols, and water, which vaporize and flow throughoverhead line 110 to a condenser 112. Most of the alcohol, water andpropylene glycol vapors condense in the condenser 112 and are collectedin the condensate tank 114 for discharge through discharge line 116 asproduct 118. Unreacted hydrogen and remaining vapors from the condenser112 are recycled back to the reactor-separator 102 through the hydrogenrecycle line 108.

Reaction products 118 are removed from the condensate tank 112 throughdischarge line 116, and the reaction mixture inside reactor-separator102 may be purged periodically or at a slow flow rate through purge line120 to obtain purge mixture 122. Purging is necessary or desirable whennon-volatile reaction by-products are formed and when metals orinorganic acids, such as residual biodiesel catalysts, are present inthe polyhydric feed stock 104. Catalysts and useful components, such asglycerol and propylene glycol, are preferably recovered from the purgemixture 122.

The reaction inside reactor-separator 102 is catalyzed, and may befacilitated at periodic intervals or by the continuous introduction of asuitable catalyst 124, which may be any catalyst that is suitable foruse in converting glycerol into lower alcohols, such as acetol and/orpropylene glycol. The catalyst 124 may reside within thereactor-separator as a packed bed, or distribution of the catalyst 124inside reactor-separator 102 may be improved by using the hydrogen gas108 to provide a fluidized bed, or by stirring (not shown). Agitatedslurry reactors of a liquid phase reaction with a vapor overhead productare preferred. The catalyst 124 is mixed with the polyhydric feedstock104 that is undergoing reaction in the reactor separator 102 tofacilitate breaking of carbon-oxygen or carbon-carbon bonds includingbut not limited to hydrogenation. As used herein, hydrogenolysis andhydrogenation are interchangeable terms.

By way of example the reaction of glycerol with hydrogen to formpropylene glycol and water is referred to frequently as hydrogenation inthis text. Suitable catalysts for this purpose may include, withoutlimitation, such metals as platinum, palladium, ruthenium, chromium,nickel, copper, zinc, rhodium, chromium, ruthenium, and combinationsthereof. Catalysts may be deposited on a substrate, such as an aluminasubstrate. In a broader sense, suitable catalysts may include thosecatalyst containing one or more elements of the subgroups from Group I,Group VI, and/or Group VIII of the Periodic Table. The best catalystsare non-volatile, and are preferably prevented from exiting the reactorseparator 102 into the condensate tank 114. A filter 125 in the overheaddischarge line 110 from the reactor separator 102 retains solidcatalysts in the reactor separator 102. No limitations are placed orimplied on whether the catalyst is soluble or solid, the oxidative stateof the catalyst, or the use of solid supports or soluble chelates.

Reaction times at preferred conditions may range from a few minutes to96 hours. Reaction time may be defined as the volume of liquid in thereactor divided by the time-averaged flow rate of liquids into thereactor. While the preferred reaction times are greater than 2 hours,the average residence time at higher loadings of catalyst 124 can beless than an hour and typically longer than 0.5 hours. While preferredtemperatures are up to 250° C., the reactor-separator may be operated attemperatures up to 270° C. with satisfactory results.

The polyhydric feed stock 104 preferably contains glycerol. In a broadersense, polyhydric feedstock 104 may contain, for example, from 5% tosubstantially 100% of a polyol, for example, glycerol, sorbitol,6-carbon sugars, 12-carbon sugars, starches and/or cellulose.

As illustrated in FIG. 1, the process equipment 100 is preferablyconfigured to provide hydrogen 106 as a reagent; however, the use ofhydrogen is optional. Commercially valuable products may be formed asintermediates that collect in the condensate tank in the absence ofhydrogen. Accordingly, use of hydrogen 106 is preferred, but notnecessary. For example, the intermediates collecting in condensate tank114 may include acetol (hydroxy-2-propanone), which may be subjected tohydrogenolysis by at least two mechanisms as shown in FIGS. 2 and 3. Inaddition to reagents, the material within reactor separator 102 maycontain water, salts, or catalysts residue from previous processes.

One type of polyhydric feedstock 104 may contain glycerol that isprepared by transesterification of oils or fatty acids, for example, asdescribed in co-pending application Ser. No. 10/420,047 filed Apr. 23,2003, which is incorporated by reference to the same extent as thoughfully replicated herein. In a polyhydric feedstock 104 of this type,water may be present in an amount ranging from 0% to 70%. Morepreferably, water is present in an amount ranging from 5% to 15%. Watermay be added to reduce side-reactions, such as the formation ofoligomers.

One advantage of using the process equipment 100 is that volatilealcohol products are removed from the reaction mixture as they areformed inside reactor separator 102. The possibility of degrading theseproducts by continuing exposure to the reaction conditions iscommensurately decreased by virtue of this removal. In addition, thevolatile reaction products are inherently removed from the catalysts toprovide relatively clean products. This reaction-separation technique isespecially advantageous for catalysts that are soluble with oremulsified in the reaction mixture.

A preferred class of catalyst 124 is the copper chromite catalyst,(CuO)_(x) (Cr2O3)_(y). This type of catalyst is useful in the processand is generally available commercially. In this class of catalyst, thenominal compositions of copper expressed as CuO and chromium expressedas Cr₂O₃ may vary from about 30-80 wt % of CuO and 20-60 wt % of Cr₂O₃.Catalyst compositions containing about 40-60 wt % copper and 40-50 wt %of chromium are preferred.

Preferred catalysts for use as catalyst 124, in addition to the copperand chromium previously described, also include barium oxide andmanganese oxide or any of their combinations. Use of barium andmanganese is known to increase the stability of the catalyst, i.e., theeffective catalyst life. The nominal compositions for barium expressedas barium oxide can vary 0-20 wt % and that for manganese expressed asmanganese oxide can vary from 0-10 wt %. The most preferred catalystcompositions comprise from 40%-60 wt % of CuO 40-55 wt % of Cr₂O₃, 0-10wt % of barium oxide and 0-5 wt % manganese oxide.

Reaction Mechanism

According to one mechanism proposed by Montassier et al. (1988),dehydrogenation of glycerol on copper can form glyceric aldehyde inequilibrium with its enolic tautomer. The formation of propylene glycolwas explained by a nucleophilic reaction of water or adsorbed OHspecies, a dehydroxylation reaction, followed by hydrogenation of theintermediate unsaturated aldehyde. This reaction mechanism was observednot to apply in our investigation.

FIG. 2 shows a preferred reaction mechanism 200 for use in thereactor-separator 102 of FIG. 1, and for which process conditions may besuitably adjusted as described above. As shown in FIG. 2, hydroxyacetone(acetol) 202 is formed, and this is possibly an intermediate of analternative path for forming propylene glycol by a different mechanism.The acetol 202 is formed by dehydration 204 of a glycerol molecule 206that undergoes intramolecular rearrangements as shown. In a subsequenthydrogenation step 208, the acetol 202 further reacts with hydrogen toform propylene glycol 210 with one mole of water by-product resultingfrom the dehydration step 204.

Early studies to investigate the effect of water on the hydrogenolysisreaction indicated that the reaction takes place even in absence ofwater with a 49.7% yield of propylene glycol. Moreover, and by way ofexample, the reaction is facilitated by use of a copper-chromitecatalyst, which may be reduced in a stream of hydrogen prior to thereaction. In this case, the incidence of surface hydroxyl species takingpart in the reaction is eliminated. The above observations contradictthe mechanism proposed by Montassier et al. where water is present inthe form of surface hydroxyl species or as a part of reactants.

EXAMPLE 1 Confirmation of Reaction Mechanism

An experiment was performed to validate the reaction mechanism 200.Reactions were conducted in two steps, namely, Steps 1 and 2. In step 1,relatively pure acetol was isolated from glycerol. Temperature rangedfrom 150° C. to 250° C. and more specifically from 180° C. to 220° C.There was an absence of hydrogen. Pressure ranged from 1 to 14 psi (6.9MPa to 96 MPa) more specifically from 5 to 10 psi (34 MPa to 69 MPa). Acopper-chromite catalyst was present. In Step 2, the acetol formed inStep 1 was further reacted in presence of hydrogen to form propyleneglycol at a temperature ranging from 150° C. to 250° C. and morepreferably between 180 to 220° C. Excess hydrogen was added at ahydrogen over pressure between 1 to 25 bars using the same catalyst.

It was observed in the Step 2 of converting acetol to propylene glycolthat lactaldehyde was formed. Propylene glycol is also formed by thehydrogenation 208 of lactaldehyde 302, as illustrated in FIG. 3. Withrespect to FIG. 2, lactaldehyde represents an alternative path forforming propylene glycol from acetol. FIG. 3 shows this mechanism 300where the acetol undergoes a rearrangement of the oxygen double bond toform lactaldehyde 302, but the dehydrogenation step 208 acting upon thelactaldehyde 302 also results in the formation of propylene glycol 210.It was also observed that the formation of lactaldehyde intermediate waspredominant at lower reaction temperatures in the range of from 50° C.to 150° C. (see Example 8 below).

The embodiments of this disclosure include production of lactaldehyde. Aprocess for converting glycerol to lactaldehyde with high selectivitypreferably includes the steps of combining a glycerol-containingfeedstock with less than 50% by weight water with a catalyst that iscapable of dehydrating glycerol to form a reaction mixture; and heatingthe reaction mixture to a temperature between 150° C. to 200° C. over areaction time interval between 0 to 24 hours at a pressure between 0.02and 25 bar. Preferably the catalyst used in the step of combining iscontains an element of the subgroups from Group I, Group VI, and/orGroup VIII of the Periodic Table. Preferably the glycerol-containingfeedstock used in the step of combining contains from 0% to 15% water inglycerol by weight. Preferably the catalyst used in the step ofcombining is a heterogeneous catalyst selected from the group consistingof palladium, nickel, rhodium, copper, zinc, chromium and combinationsthereof. Preferably the process includes a step of removing reactionproduct vapors that form during the step of heating. Preferably theprocess includes a step of condensing the vapors to yield liquidreaction product. Preferably temperature used in the heating step existswithin a range from 165° C. to 185° C. Preferably the pressure used inthe heating step exists within a range from 0.02 to 2 bars. Preferably,the step of condensing occurs using a condenser operating at atemperature between 0° C. to 140° C.

This and subsequent reactions were performed in liquid phases withcatalyst and sufficient agitation to create a slurry reaction mixture.

EXAMPLE 2 Simultaneous Dehydration and Hydrogenation Using VariousCatalysts and Reagent Mixtures

A variety of reaction procedures were performed to show that reactionefficiency may be optimized at any process conditions, such as reactiontime, temperature, pressure and flash condition by the selection orchoice of catalyst for a given polyhydric feedstock.

Table 1 reports the results of reacting glycerol in the presence ofhydrogen and catalyst to form a mixture containing propylene glycol. Thereaction vessel contained 80 grams of refined glycerol, 20 grams ofwater, 10 grams of catalyst, and a hydrogen overpressure of 200 psig.The reactor was a closed reactor that was topped off with excesshydrogen. The reaction occurred for 24 hours at a temperature of 200° C.All catalysts used in this Example were purchased on commercial orderand used in the condition in which they arrived.

TABLE 1 Summary of catalyst performances based on 80 grams of glycerolreported on a 100 grams basis. Catalyst Best 5% Catalyst Initial Pos-Ruthenium Catalyst Raney- Loading sible on carbon Raney- Nickel (g) (g)(g) Copper (g) (g) Glycerol 100 0 63.2 20.6 53.6 Water 25 43 not not notmeasured measured measured Propylene Glycol 0 82 14.9 27.5 14.9 EthyleneGlycol 0 0 16.9 13.1 16.5 Acetol 0 0 0.0 12.1 0.0 Total, excluding 10082 94.9 73.2 85.0 water

Table 2 summarizes reaction performance with a higher initial watercontent, namely, 30 grams of refined glycerol and 70 grams of water. Thereactions were conducted at the following initial conditions: 5% wt ofcatalyst, and a hydrogen overpressure of 1400 kPa. The following tablepresents compositions after reacting in a closed reactor (with toppingoff of hydrogen) for 24 hours at a reaction temperature of 200° C.

TABLE 2 Summary of catalyst performances based on 30 grams initialloading of glycerol and 70 grams of water. Catalyst Catalyst InitialBest Catalyst 5% Raney- Raney- Loading Possible Ruthenium Copper Nickel(g) (g) on carbon (g) (g) (g) Glycerol 30 0 20.8 19.1 3.8 PropyleneGlycol 0 24 9.3 7.23 3.1 Ethylene Glycol 0 0 0 0 0 Acetol 0 0 1.5 1.61.7

Table 3 summarizes the performance of a copper chromium catalyst in thepresence of 20 percent of water. The reactions were conducted at thefollowing initial conditions: 5% wt of catalyst, and a hydrogenoverpressure of 1400 kPa. The following table presents compositionsafter reacting in a closed reactor (with topping off of hydrogen) for 24hours at a reaction temperature of 200° C.

TABLE 3 Summary of copper chromium catalyst performance based on 80grams initial loading of glycerol and 20 grams of water. Initial BestCatalyst Copper Loading (g) Possible (g) Chromium (g) Glycerol 80 0 33.1Propylene glycol 0 66.1 44.8 Ethylene Glycol 0 0 0 Acetol 0 0 3.2

Table 4 summarizes the impact of initial water content in the reactantson formation of propylene glycol from glycerol. The reactions wereconducted at the following initial conditions: 5% wt of catalyst, and ahydrogen overpressure of 1400 kPa. The catalyst was purchased fromSud-Chemie as a powder catalyst having 30 m²/g surface area, 45% CuO,47% Cr₂O₃, 3.5% MnO₂ and 2.7% BaO. The following table presentscompositions after reacting in a closed reactor (with topping off ofhydrogen) for 24 hours at a reaction temperature of 200° C.

TABLE 4 Summary of catalyst performances based on different initialloadings of glycerol in water. Water (wt %) % Conversion % Yield %Selectivity 80 33.5 21.7 64.8 40 48 28.5 59.4 20 54.8 46.6 85.0 10 58.847.2 80.3 0 69.1 49.7 71.9

The reaction was performed using a small scale reaction distillationsystem like that shown as process equipment 100 in FIG. 1 to process areaction mixture including 46.5 grams of refined glycerol and 53.5 gramswater. The catalyst was purchased from Sud-Chemie as a powder catalysthaving 30 m²/g surface area, 45% CuO, 47% Cr₂O₃, 3.5% MnO₂ and 2.7% BaO.Table 5 summarizes performance with higher initial water content using asmall reaction distillation system.

TABLE 5 Example of reaction distillation. Reactor Distillate Glycerol21.6 grams 2.2 Propane Diol 6.4 9.5 Ethylene Glycol 0 0 Acetol 1.4 1.4Use of Glycerol from Fatty Acid Glyceride Refinery

One preferred source of the polyhydric feedstock 104 is crude naturalglycerol byproducts or intermediates, for example, as may be obtainedfrom processes that make or refine fatty acid glycerides frombio-renewable resources. These are particularly preferred feedstocks formaking an antifreeze mixture. When using these feedstocks, theantifreeze mixture is prepared as explained above by hydrogenation ofglycerol over a catalyst, which is preferably a heterogeneous catalyst.The reactor-separator 102 may, for example, be a packed bed reactor,slurry, stirred or fluidized bed reactor. When the hydrogenationreaction is performed in a packed-bed reactor, the reactor effluent islargely free of catalyst. In the case of a slurry reactor, aheterogeneous catalyst may be filtered from the reactor effluent. Thereactor-separator 102 may be used for slurry reactions by circulatinghydrogen from the top vapor phase to the bottom of the reactor to createincreased agitation and by preferably using a catalyst that has adensity similar to the density of the liquid in the reactor. A fluidizedbed may be used where the densities differ, where a catalyst bed isfluidized by the incoming hydrogen from line 108. Conventional agitationmay also promote hydrogen contact in the liquid.

To make antifreeze, the process conditions need only provide moderatehydrogenation conversions of glycerol, e.g., those ranging from 60% to90% conversion. This is because from 0% to 40% of the glycerol in thepolyhydric feedstock 104 on a water-free basis may remain with propyleneglycol products in the antifreeze product. For some productapplications, the final antifreeze product may suitably contain up to60% glycerol. Furthermore, when the product 118 contains a low glycerolconcentration, e.g., less than 40% where there is an effectiveconversion of 60% to 90%, other known antifreezes may be mixed with theproducts 118. Alternatively, the purge materials 122 may be mixed withthe contents of condensate tank 114, for example, after filtering, toform a salable product that may be directly discharged from the processequipment 100.

One particularly preferred source of polyhydric feedstock 104 for thereaction is the natural glycerol byproduct that is produced during thevalue-added processing of naturally occurring renewable fats and oils.For example, the glycerol byproduct may be a vegetable oil derivative,such as a soy oil derivative. This variety of polyhydric feedstock 104may contain water, soluble catalysts, and other organic matter that arepresent in intermediate mixtures which are produced in the manufactureof glycerol for sale into the glycerol market. One advantage of thepresent instrumentalities is that little or no refining of theseintermediates are necessary for their use as polyhydric feedstock 104 inmaking commercial antifreeze or deicing mixtures.

These intermediates and other polyhydric feedstocks 104 may contain highamounts of water. The ability to use polyhydric feedstocks 104 thatcontain high amounts of water advantageously reduces costs for thisprocess over other uses for the glycerol. The water content both in thepolyhydric feedstock 104 prior to the reaction and in the salablereaction product is generally between 0 and 50%.

The polyhydric feedstock 104 may contain residual catalyst that wasadded during alcoholysis of these intermediates. The fate of solubleresidual catalysts, i.e., those that remain from alcoholysis in thepolyhydric feedstock 104 and which are in the purge material 122 dependsupon:

-   -   1. the specific type of soluble residual catalyst, and    -   2. any interaction between the residual catalyst and another        catalyst that is added to the crude glycerol to promote        hydrogenation within reactor-separator 102.

The residual catalyst content in the glycerol feedstock 104 from theprocessing of bio-renewable fats and oils is commonly between 0% and 4%or even up to 10% by weight on a water-free basis. One way to reduce theresidual catalyst content is to minimize the amount that is initiallyused in alcoholysis of the fatty acid glyceride. The alcoholysis may,for example, be acid-catalyzed. Neutralizing the residual catalyst withan appropriate counter-ion to create a salt species that is compatiblewith the antifreeze specifications is preferred to removing the residualcatalyst.

Alternatively, neutralization can be performed to precipitate thecatalyst from the liquid glycerol. Calcium-containing base or salt maybe used to neutralize the residual catalyst in the polyhydric feedstock104, and the solid salts generated from this neutralization may beseparated from the liquid, for example, by filtration or centrifugationof effluent from reactor-separator 102, such as by filtering purgematerial 122. Acid-base neutralization to form soluble or insolublesalts is also an acceptable method of facilitating separation.Specifically, neutralizing potassium hydroxide with sulfuric acid toform the dibasic salt is a acceptable procedure. As shown by way ofexample in FIG. 1, neutralization of sodium or potassium catalyst, whichis sometimes introduced into the value-added processing method for fatsand oils, can be achieved by adding stoichiometric equivalent amounts ofa neutralizing agent 126, such as calcium oxide and/or sulfuric acid, toform the calcium salt of the catalyst. These salts are largely insolubleand may be filtered from the purge material 122. To improve separationof the substantially insoluble salt, the water content is preferablyreduced to less than 20% by weight and the filtration is preferablyperformed at temperatures less than 40° C. and more preferably below 30°C. The optimal filtration temperature depends upon composition where thereduced solubility of salts at lower temperatures is weighed againstlower viscosities at higher temperatures to identify the best filtrationconditions.

One general embodiment for processing of crude glycerol to antifreeze inthe fatty acid glyceride refinery embodiment follows a C₁ to C₄ alkylalcohol alcoholysis process. The incoming crude glycerol feedstock 104is neutralized by the addition of a neutralizing agent 126 to achieve apH between 5 and 12, which is more preferably a pH between 5 and 9. TheC₁ to C₄ alcohol and water are separated by distillation from the crudeglycerol, such that the combined concentrations of water and C₁ to C₄alcohols within reactor-separator 102 are less than 20 wt % by weightand, preferably, less than 5% by weight. In a stepwise process where thepolyhydric feedstock 104 is added to the reactor-separator 104 atperiodic intervals, selected components of these alcohols and/or theirreaction products may be isolated by fractional distillation throughoverhead line 110 and discharged from condensate tank 114. This may bedone by flash liberation of such alcohols at suitable times to avoid orlimit their combining with propanediols, according to the principle offractional distillation. Subsequent hydrogenation of the flashedglycerol within reactor-separator 102 suitably occurs by contacting thecrude glycerol with a hydrogenation catalyst and hydrogen at a pressureranging from 1 bar to 200 bar and at a temperature ranging from 100° to290° C. until a conversion of the glycerol between 60% and 90% isachieved. More preferably, process conditions entail the contactpressure for hydrogenation ranging from 1 to 20 bar.

Separating the C₁ to C₄ alcohol and water is preferably achieved byselective flash separation at temperatures greater than 60° C. and lessthan 300° C. Alternatively, separating the C₁ to C₄ alcohol and watermay be achieved in a process based on thermal diffusion, as is describedin related application Ser. No. 10/420,047, where for example thereactor-separator 102 is a thermal diffusion reactor. Alternatively,water is added prior to hydrogenation as water promotes hydrogenation inthe presence of certain catalysts.

The amount of organic matter in the polyhydric feedstock issubstantially dependent upon the fat or oil from which the glycerol wasobtained. The organic matter (other than glycerol) is typically fattyacid derivatives. One method for mitigating residual organic matter isby filtration. Alternatively, it is possible to decant insolubleorganics from the glycerol in a gravity separator (not shown) attemperatures between 25 and 150° C. As necessary, the flash point of themixture is preferably increased to greater than 100° C. by flashseparation of volatiles from the glycerol-water mixture. Specifically,the residual C₁ to C₄ alkyl alcohol content in the feedstock is flashliberated to achieve feedstock concentrations that are preferably lessthan 1% alkyl alcohol. Depending upon the alkyl alcohol, vacuum may needto be applied to reach achieve the 1% alkyl alcohol concentration.

The following are preferred reaction conditions for conversion for usein processing these feedstocks. These are similar but not exactly thesame as preferred+conditions that have been previously described for usein the reactor-separator 102. The reaction temperature is 150° C. to250° C. The reaction time is from 4 to 28 hours. Heterogeneous catalystsare used which are known to be effective for hydrogenation, such aspalladium, nickel, ruthenium, copper, copper zinc, copper chromium andothers known in the art. Reaction pressure is from 1 to 20 bar, buthigher pressures also work. Water in the polyhydric feedstock ispreferably from 0% to 50% by weight, and more preferably from 5 to 15%water by weight.

The preferred reaction conditions provide a number of performanceadvantages. Operating at temperatures less than 250° C. dramaticallyreduces the amount of unintended by-product formation, for example,where lower concentrations of water may be used without formation ofpolymers or oligomers. Furthermore, operation at temperatures near 200°C., as compared to near 300° C., provides an increased relativevolatility of propylene glycol that facilitates an improved separationof propylene glycol from the glycerol reaction mixture. The use of lowerpressures allows the use of less expensive reaction vessels, forexample, as compared to high-pressure vessels that operate above about28 bars, while also permitting the propylene glycol to distill fromsolution at these temperatures. Even so, some embodiments are notlimited to use at pressures less than 20 bars, and may in fact bepracticed at very high hydrogen pressures. The disclosed processconditions are viable at lower pressures (less than 20 bar) whereas mostother processes to produce similar products require much higherpressures.

By these instrumentalities, glycerol may also be hydrogenolysed to 1,2and 1,3 propanediols. The 1,3 propanediol may be optionally separatedfrom this mixture by methods known in the science and used as a monomerwhile the remaining glycerol and propanediols are preferably used asantifreeze.

EXAMPLE 3 Packed-Bed Reactor Embodiments

One method of preparing acetol and propylene glycol from glycerolincludes a gas phase reaction at a temperature ranging from 150° to 280°C. in a packed-bed reactor. In some embodiments this temperature is morepreferably from 180° C. to 240° C. or 250° C. to avoid thermaldegradation of reaction products. The reactions described hereinoccurred in a packed-bed reactor. The pressures in the reaction vesselare preferably from 0.02 to 25 bars and in some embodiments thispressure is more-preferably between 0.02 and 10 bars. Most preferably,the reaction pressure exists within a range from 0.2 and 1.2 bars.

FIG. 4 provides a block flow diagram of process equipment 400 includingan evaporator 426 for creating a vapor reaction mixture 427.Non-volatile components 428 in the polyhydric feed 104 are removed fromthe evaporator 426 in a continuous or semi-batch mode. The evaporator426 is particularly effective for processing crude glycerol thatcontains salts where, otherwise, the salts poison the catalyst. Apolyhydric feed stock 104, for example, containing glycerol, isintroduced stepwise or continuously into the evaporator 426. The vaporreaction mixture 427 proceeds to the packed-bed reactor 425 where theheterogeneous catalyst promotes conversion of glycerol 104 to acetol andpropylene glycol in sequential reactions. The vapor product mixture 430proceeds to the condenser 431 where a condensate product is formed 432and proceeds to product storage 433. The gas effluent may disposedthrough purge or vacuum 434.

Water is produced as a reaction byproduct and may be kept with thepropylene glycol product or removed. A major advantage of the currentprocess over other processes in the literature is the very lowconcentration or absence of ethylene glycol resultant from either theuse of copper chromite catalyst or formation and purification of acetolas an intermediate. The acetol can be readily purified from any ethyleneglycol prior to hydrogenation by distillation.

The processes of this operation may be maintained at pressures below 1bar through the use of a vacuum source preferably connected to thecondensation process at the end of the process. In the most ideal ofcases, the condenser 431 itself can maintain pressures less than 1 bar;however, from a practical perspective, a vacuum is needed to pull offany inert gases (nitrogen etc.) that may accumulate in the system.

The reaction system of FIG. 4 is effective for producing either acetolor propylene glycol. FIG. 5 provides a schematic diagram of thelaboratory equipment 500 showing a variation of this equipment 500,which includes an evaporator 526 connected to the packed bed reactor525. Vapor effluent from packed bed reactor 525 is condensed in thecondenser 531 by action of a cold bath to draw heat 538. Heat 540 isapplied to the evaporator 526 to create the vapor reagent. A vacuum 542connected to the condenser 531 literally pulls the vapors through thesystem and allows the glycerol feed to evaporate at lower temperaturesthan would occur at atmospheric pressure. An oil bath 534 maintains thepacked bed reactor 535 at a predetermined temperature or temperaturerange by flow of heat 542. The glycerol is loaded into the evaporator526 at the start of the experiment and may be added through an auxiliaryfeed to the evaporator (not shown) during the experiment. An optionalgas feed 544 that contains nitrogen and may also contain hydrogen isdirected to the evaporator 526.

The process equipment shown in FIG. 5 was used to react glycerol undervarious conditions. Various runs were made using the equipment 500according to the materials and conditions reported in Table 6. ReactionG1 of Table 6 provides example conversion data over 3.3 mm cylindricalpellets of copper chromite catalyst. The pressure of this reaction wasless than 0.1 bar, and the temperature is about 230° C. Reaction G1illustrates the effectiveness of the gas phase reaction over apacked-bed of catalyst for producing acetol in high selectivity.

TABLE 6 Summary of gas phase reactor performances in packed-bed reactor.Area of by-product Acetol PG Glycerin Acetol + PG Water 10.77/Area ofTotal Mass RXN/Date Conditions Sample (wt %) (wt %) (wt %) (wt %) (wt %)standard balance (wt %) G1 Catalyst packing 1 13.44 0.97 72.92 14.418.25 0.12 95.58 Before (size: 3 * 3 mm) ~50-60 g 2 13.43 1.08 77.8814.51 8.07 0.11 100.46 2005/9/29 Control (No gas Purge) 3 12.24 1.0669.51 13.3 12 0.12 94.81 Reactor temp: 230 C. Pressure: 29.9 in-Hg (vac)Proof of Concept—Low pressure dehydration reaction works. G2 Catalystpacking 1 19.01 1.68 70.70 20.69 5.86 0.10 97.25 Before (size: 3 * 3 mm)~50-60 g 2 18.42 1.79 72.42 20.21 5.38 0.11 98.01 2005/9/29 HydrogenPurge 3 16.24 1.56 75.50 17.80 5.21 0.10 98.51 Reactor temp: 230 C. 415.81 1.62 75.87 17.43 4.97 0.11 98.27 Pressure: 26 in-Hg (vac)Experiment demonstrates that hydrogen partial pressure reduces waterformation and leads to improved mass balance-better yields. Conversionsappeared to be higher. G3 Catalyst packing 1 9.75 0.53 87.27 10.28 3.870.16 101.42 Before (size: 3 * 3 mm) ~50-60 g 2005/9/29 Nitrogen PurgeReactor temp: 230 C. Pressure: 26 in-Hg (vac) Nitrogen was not as goodas hydrogen based on higher water content of nitrogen reaction.Theoretical water is 1 part water for four parts acetol (acetol +propylene glycol). Actual water is greater than theoretical. The ratioof the by-product peak to desired product is higher for this nitrogenrun. G4 Catalyst packing 2005/9/29 (size: 3 * 3 mm) 50 g HOT PLATE (Nogas Purge) 1 9.88 2.41 79.15 12.29 9.05 0.12 100.49 Reactor temp: 230 C.2 12.65 0 77.26 12.65 8.03 0.14 97.94 Pressure: 27 in-Hg (vac) This runsummarizes a different feed mechanism where feed is put on a hot plateto evaporate feed as it is introduced. Method provided improvedexperimental control but did not lead to new insight into the reaction.G5 Catalyst packing 2005/10/4 (size: 3 * 3 mm) 50 g Hydrogen Purge 112.52 2.86 78.4 15.38 4.18 0.11 97.96 Nitrogen Purge 1 7.79 0.81 87.948.60 3.56 0.14 100.1 Reactor temp: 230 C. Pressure: 27 in-Hg (vac) Theseruns are a repeat comparison of the use of hydrogen versus nitrogen. Thehydrogen provided higher yields, more propylene glycol, less additionalwater, and fewer junk peaks. Motivation for increased use of hydrogenwas the fat that production of PG must grab a hydrogen from somewhere,and that somewhere could only be other glycerin or acetolproducts-leading to the hypothesis that addition of hydrogen wouldincrease the yield of desired products. G6 Catalyst packing (size: 9-401 23.05 1.66 73.1 24.71 5.00 0.07 102.81 2005/10/5 mesh) 50 g- fresh 224.3 1.49 72.21 25.79 5.07 0.08 103.07 Hydrogen Purge Reactor temp: 230C. Pressure: 27 in-Hg (vac) This run summarizes the impact of usingsmaller catalyst. The conversion increased by 50%. G7 Catalyst packing(size: 9-40 1 44.74 2.4 37.49 47.14 12.00 0.06 96.63 2005/10/10 mesh)100 g- fresh 2 42.56 2.3 38.01 44.86 9.93 0.07 92.8 Hydrogen PurgeReactor temp: 230 C. Pressure: 27 in-Hg (vac) This run summarizes theimpact of using more smaller catalyst. Doubling the catalystconcentration doubled the conversion. To a first approximation, thisreaction is zero-order. G8 Catalyst packing (size: 9-40 1 64.11 6.42 4.370.53 19.00 0.00 93.83 2005/10/11 mesh) 150 g-100 g used 1, 2 63.14 5.644.46 68.78 19.25 0.00 92.49 50 g fresh 3 63.73 5.28 7.28 69.01 17.880.00 94.17 Hydrogen Purge Reactor temp: 240 C. Oil batch temp: 232 C.Pressure: 27 in-Hg (vac) This run summarizes the impact of using evenmore smaller catalyst smaller catalyst. Tripling the catalystconcentration (50 to 150 grams) tripled the conversion. To a firstapproximation, this reaction is zero-order. Total glycerin reacted:369.11 g Total products: 360.52 g Reaction time: 2.5 hr Mass balance ofglycerin In versus product formed is pretty good for this system. G9Catalyst packing (size: 9-40 1 62.35 7.51 6.25 69.86 18.57 0.00 94.682005/10/11 mesh) 150 g-100 g used 2, 2 64.28 5.03 7.24 69.31 18.89 0.0095.44 50 g used 1 3 60.39 4.39 14.11 64.78 19.01 0.06 97.9 HydrogenPurge Reactor temp: 240 C. Oil batch temp: 232 C. Pressure: 27 in-Hg(vac) Total glycerin reacted: 754.79 g Total products: 750.40 g Reactiontime: 5 hr This extended run shows good mass balance of glycerin inversus product out. Slight decrease in conversion with time deemed to bewithin experimental error. G10 Acetol to PG with H2 Purge on a HOT PLATE2005/10/14 Catalyst packing (size: 9-40 mesh) 150 g Pressure: 27 in-Hg(vac), 1 53.4 22.53 5.42 81.35 closed valve Pressure: 20 in-Hg (vac), 242.16 11.09 37.38 90.63 open valve to maintain pressure Pressure: 20in-Hg (vac), 3 33.57 14.29 43.94 91.8 more H2 flow to maintain pressureImproved Packed-Bed Embodiments

It was observed that propylene glycol was produced in illustrativeexample G1 of Table 6. Since the only source of hydrogen for reactingwith acetol (or glycerol) to form propylene glycol was from anotheracetol or glycerol molecule it was hypothesized that the absence of freehydrogen in the system led to scavenging of hydrogen from the glyceroland that this scavenging led to undesired byproducts and loss in yield.

To overcome the problem with scavenging of hydrogen from glycerol, asmall amount of hydrogen was introduced to the system. FIG. 6illustrates the preferred packed-bed reaction process equipment 600 withhydrogen feed 636 as a modification to the process equipment of FIG. 4The hydrogen feed 636 was introduced to the evaporator 426 since thisgas diluents, in addition to being useful in hydrogenating acetol, alsopromotes evaporation of glycerol. Since glycerol has a vapor pressure ofa mere 0.15 bar at 230° C., the hydrogen overpressure can add to thispressure to increase overall pressure—but this is primarily possible ifglycerol is evaporated in the presence of a gas like hydrogen. Thecondenser 431 condenses the acetol and propylene glycol from unreactedgas. although the unreacted gas may be purged 434, a recycle loop 640may be used to resupply the evaporator 426, packed bed reactor 425, orcondenser 431 by selective arrangement of valves 642, 644, 646.

Reaction G2 of Table 6 provides example conversion data illustrating thebeneficial impact of a hydrogen feed (purge) 544 (see FIG. 5) as thehydrogen feed 636 of FIG. 6 combined with the glycerol feed in thepacked-bed reactor 425. The pressure of was 0.13 bars, and thetemperature was 230° C. A higher yield to acetol and propylene glycolwas observed.

The desired dehydration reactions produce one water molecule for everyacetol (or propylene glycol) molecule that is formed. Water present inexcess of this indicates excess dehydration and lower selectivities. Theratio of actual to theoretical water content decreased from 2.3-3.6 to1.07-1.17 as a result of hydrogen being present during the dehydrationreaction. In addition, a GC peak at 10.77 minutes is a by-product. Theratio of this peak area to the mass fractions of desired acetol andpropylene glycol decreased from 0.76-0.9 to 0.47-0.63 as a result ofhydrogen being present during the dehydration reaction.

To confirm that the desired results were a result of hydrogen ratherthan any diluent in the system, experiment G3 was performed usingnitrogen instead of hydrogen. The ratio of actual to theoretic waterincreased to 1.51 with nitrogen. In addition the ratio of the 10.77minute peak increased to 1.56.

Both the hydrogen and nitrogen diluent/purge experiments were repeatedin experiment G5 with generally repeatable results and validation of thebenefit of using hydrogen as a diluent/purge during the dehydrationreaction that forms primarily acetol as a product.

The preferred process uses a hydrogen diluent and reagent 636 introducedto the evaporator 426.

The following are a summary of the experiments summarized in Table 6 andwhat the results indicate:

-   -   Experiment G1 provides proof of concept of low-pressure        dehydration over a packed-bed catalyst.    -   Experiment G2 demonstrates that hydrogen partial pressure        reduces water formation and leads to improved mass        balance—better yields. Conversions appeared to be higher.    -   Experiment G3 demonstrates that nitrogen was not as good as        hydrogen based on higher water content of nitrogen reaction.        Theoretical water is 1 part water for four parts acetol        (acetol+propylene glycol). Actual water for this experiment is        considerably greater than theoretical. The ratio of the        by-product peak (10.77 minutes) to desired product is higher for        this nitrogen run.    -   Experiment G4 demonstrates a continuous feed mechanism approach        where feed is put on a hot plate to evaporate feed as it is        introduced. Method provided improved experimental control but        did not lead to new insight into the reaction.    -   Experiment G5 provides a repeat comparison of the use of        hydrogen versus nitrogen. The hydrogen provided higher yields,        more propylene glycol, less additional water, and fewer junk        peaks. Motivation for increased use of hydrogen was the fat that        production of PG must grab a hydrogen from somewhere, and that        somewhere could only be other glycerol or acetol        products—leading to the hypothesis that addition of hydrogen        would increase the yield of desired products.    -   Experiment G6 summarizes the impact of using smaller catalyst.        The conversion increased by 50%.    -   Experiment G7 summarizes the impact of using more smaller        catalyst. Doubling the catalyst mass doubled the conversion. To        a first approximation, this reaction is zero-order.    -   Experiment G8 summarizes the impact of using even more smaller        catalyst smaller catalyst. Tripling the catalyst mass (50 to 150        grams) tripled the conversion. To a first approximation, this        reaction is zero-order.    -   Experiment G9 summarizes a good mass balance of glycerol        relative to the reaction products.    -   Experiment G10 repeats the mass balance run of G9 illustrating a        good mass balance of glycerol in versus product out. A slight        decrease in conversion with time was deemed to be within        experimental error.

Experiments validated conversions of greater than 95% for the conversionof glycerin to acetol. At conversions of about 98%, approximately 70%acetol and 9% propylene glycol were present in the product. Continuedcontact of both hydrogen and acetol over the copper chromite catalystcontinued to increase yields to propylene glycol.

Other Packed-Bed Embodiments

FIG. 7 shows a more preferred process of preparing propylene glycol fromglycerol.

FIG. 7 provides a block flow diagram of process equipment 700 includingglycerol (or polyhydric) 704 and hydrogen feeds 736. The hydrogen iscontacted with the glycerol in an evaporator 726 operated between about200 and 250° C., which promotes evaporation of glycerol to form a vaporreactor enfluent 727. A first packed-bed reactor 725 converts glycerolto acetol with some formation of propylene glycol.

If hydrogen is present, the acetol will react with hydrogen to formpropylene glycol in the first packed-bed reactor 725. At low hydrogenpartial pressures, about 0.1 bar, acetol is predominantly formed. Athigher hydrogen pressures, more propylene glycol is formed. Theformation of acetol is predominantly rate limited. The reaction ofacetol to propylene glycol is fast relative to the reaction of glycerolto acetol; however, the acetol to propylene glycol reaction isequilibrium limited. Since reaction of acetol to propylene glycol isequilibrium limited, recycle of acetol is useful to maximize productionof propylene glycol. Distillation can be used to concentrate acetol fromthe product stream for recycle to the evaporator or other locationsprior to the reactor.

One method of operating the process of FIG. 7 is to add additionalhydrogen to stream 741 whereby acetol is primarily formed in the firstpacked-bed reactor 725 and propylene glycol is primarily formed in thesecond packed-bed reactor 738.

The dehydration reaction in the first packed-bed reactor 725 is highlyexothermic. For example, the heat of reaction will increase glycerolinitially at 200° C. to acetol at 414° C. at 100% conversion and withoutany solvent/diluent. The higher temperatures lead to a loss ofproduction and generation of undesired by-products. Heat must becontinuously or stepwise removed from the reaction mixture to maintaintemperatures below 250° C., preferably below 230° C., and mostpreferably below 220° C.

The mixture is preferably cooled to about 220° C. in a heat exchanger740 prior to hydrogenolysis 741 in packed-bed reactor two 738. Reactortwo 738 is preferably a packed-bed reactor. Copper chromite catalyst iseffective in reactor two; however, other hydrogenation catalysts mayalso be used such as Raney-nickel catalyst. The hydrogenolysis reactionis also highly exothermic.

Although high conversions are possible for both the dehydration andhydrogenolysis reaction, a separator 745 is used to further purify theproduct 733.

Preferably, the reactor 1 effluent 730 and reactor 2 effluent 739 arerecycled 742/743 along with the overheads 747 of the separator 745. Ablower 744 is used to overcome pressure drops for the recycle. If theseparator overhead 747 is a liquid, it is pumped rather than compressed.The hot recycle streams 742/743 may have temperatures up to 300° C. andreduce or eliminate the need for auxiliary heat addition to theevaporator 726. This direct contact heat exchange and evaporation isvery efficient. These recycle streams serve the addition purpose ofproviding additional heat capacity to the reactor enfluents 727/741, andthis additional heat capacity minimizes temperature increases in thereactors. Minimizing temperature increases maximizes yields of acetoland propylene glycol.

Recycled vapors 742/743/747 add additional partial pressures fromacetol, water, and propylene glycol; combined, these may add from 0.2 to1.2 bars of partial pressure. Recycle stream 1042 has the advantage ofproviding heat to the evaporator 1026, but has the disadvantage ofincreasing the residence time of acetol that can degrade acetol. Recyclestream has advantages associated with providing heat to the evaporator.Recycle stream 1047 can be enriched in hydrogen as a recycle, which isadvantageous for the reaction, but is not advantageous for providingheat to the evaporator. The corresponding total pressure in theevaporator and downstream unit operations is about 0.3 to 1.5 bars. Thepreferred pressure is about 1.1 bars such that about 0.3 bars is freshfeed glycerol and hydrogen, about 0.4 bars is recycled hydrogen, andabout 0.4 bars is recycled water vapor-heat exchangers are preferred torecover heat from the hot reaction products into the evaporator andvapors recycled to the evaporator.

Due to the exothermic nature of the both the dehydration andhydrogenolysis reactions, temperature control is most important. Thepreferred means to control temperature is to use recycled water andhydrogen as a diluent in combination with heat exchangers between afirst and second reactor. More than two reactors is optional. Inaddition, use of cold shots of propylene glycol or water in eitherreactor 1025/1038 or the heat exchanger between the reactors 1040. Forexample, cold shots of propylene glycol from the product stream 1046 canbe used to maintain temperatures less than 250° C. throughout thesystem.

Table 7 and FIG. 8 show the impact of temperature and pressure on theratio of propylene glycol to acetol in the product for a reaction at aresidence time slightly longer than is necessary to fully react all theglycerol. Table 7 summarizes the ratio of propylene glycol to acetolwhere propylene glycol was used as the feed (not glycerol in thesystem). The fact that propylene glycol reacts to form acetol fullyvalidates that this reaction is equilibrium limited. The fact that theforward (glycerol as reactant) and backward (propylene glycol isreactant) produce essentially the same ratios of propylene glycol atsimilar temperatures and pressures indicates that the acetol topropylene glycol reaction is predominantly equilibrium limited ratherthan rate limited.

TABLE 7 Effect of Temperature and Pressure on the Formation of PropyleneGlycol from Glycerol.* Pressure Log Reactor of [PG:Acetol Reactor 1000/[PG:Acetol Temperature Discharge Acetol PG Mass Temperature TemperatureMass [° C.] [bar] [wt %] [wt %] Ratio] [K] [K] Ratio] 220 1 26.00 50.841.96 493 2.03 0.29 220 1 18.58 47.29 2.55 493 2.03 0.41 238 1 20.2927.31 1.35 511 1.96 0.13 241 1 24.45 27.73 1.13 514 1.94 0.05 240 129.70 31.35 1.06 513 1.95 0.02 220 2 22.64 56.31 2.49 493 2.03 0.40 2202 17.56 63.71 3.63 493 2.03 0.56 220 2 18.34 65.75 3.59 493 2.03 0.55221 2 16.60 56.51 3.40 494 2.02 0.53 220 2 12.74 51.20 4.02 493 2.030.60 221 2 14.88 49.38 3.32 494 2.02 0.52 237 2 23.35 35.53 1.52 5101.96 0.18 236 2 19.91 40.56 2.04 509 1.96 0.31 240 2 18.85 31.80 1.69513 1.95 0.23 220 4 10.50 69.47 6.62 493 2.03 0.82 220 4 12.55 65.835.25 493 2.03 0.72 240 4 6.95 30.91 4.45 513 1.95 0.65 240 4 12.12 51.364.24 513 1.95 0.63 *For these reactions, the total pressure ispredominantly comprised of hydrogen where the molar ratio of hydrogen toalcohols is about 13:1.

FIG. 8 illustrates the data of Table 7 as pressure dependence of theglycerol to propylene glycol reaction, at temperatures of 220° C., and240° C.

TABLE 8 Effect of Temperature and Pressure on the Formation of Acetolfrom Propylene Glycol.* Pressure Log Reactor of [PG:Acetol Reactor 1000/[PG:Acetol Temperature Discharge Acetol PG Mass Temperature TemperatureMass [° C.] [bar] [wt %] [wt %] Ratio] [K] [K] Ratio] 203 1 17.36 64.313.70 476.15 2.10 0.57 239 1 34.27 39.16 1.14 512.15 1.95 0.06 202 121.33 72.40 3.39 475.15 2.10 0.53 237 1 34.56 34.67 1.00 510.15 1.960.00 177 2 6.39 88.90 13.91 450.15 2.22 1.14 181 2 11.07 85.55 7.73454.15 2.20 0.89 184 2 11.06 87.84 7.94 457.15 2.19 0.90 181 2 11.687.06 7.51 454.15 2.20 0.88 182 2 11.03 89.24 8.09 455.15 2.20 0.91 1832 10.15 92.22 9.09 456.15 2.19 0.96 207 2 15.58 75.00 4.81 480.15 2.080.68 220 2 18.19 63.25 3.48 493.15 2.03 0.54 216 2 17.25 61.20 3.55489.15 2.04 0.55 237 2 35.1 36.08 1.03 510.15 1.96 0.01 240 2 23.6143.64 1.85 513.15 1.95 0.27 242 2 21.77 36.54 1.68 515.15 1.94 0.22 2044 10.35 80.10 7.74 477.15 2.10 0.89 239 4 21.91 58.34 2.66 512.15 1.950.43 197 4 5.8 94.01 16.21 470.15 2.13 1.21 242 4 11.29 34.45 3.05515.15 1.94 0.48 242 4 20.96 52.45 2.50 515.15 1.94 0.40 241 4 9.1541.65 4.55 514.15 1.94 0.66 240 4 13.02 59.51 4.57 513.15 1.95 0.66 2404 13.63 55.59 4.08 513.15 1.95 0.61 *For these reactions, the totalpressure is predominantly comprised of hydrogen where the molar ratio ofhydrogen to alcohols is about 13:1.

Since the conversion of acetol to propylene glycol is equilibriumlimited, reaction residence times that are longer than it takes to reactthe glycerol are not advantageous in forming more propylene glycol. Infact, at longer residence times, the product concentrations willdecrease as the acetol and/or propylene glycol continue to react to formundesired by-products. As the data in Tables 7 and 8 indicate, higherpressures (4 bar rather than 1 bar) and lower temperatures (220° C.rather than 240° C.) tend to favor formation of propylene glycol. Thesetrends are fully consistent with the exothermic nature of thehydrogenation of acetol to form propylene glycol and the fact that thisreaction results in a reduction in the total moles in the system (twomoles, one each of hydrogen and acetol, react to form one mole ofpropylene glycol).

Table 9 summarizes conversion data on a 15 foot reactor loaded with 760and 1160 grams of 3×3 mm pellet copper chromium catalyst. The reactionswere evaluated with glycerol evaporated at 230° C. The by controllingthe vacuum at the exit from the cold trap (condenser), the pressure ofthe system was able to be operated at 27, 19, and 8 inches of mercury invacuum (0.074, 0.34, and 0.71 bars absolute pressure). This increase inpressure causes the partial pressure and the stoichiometric excess ofhydrogen to increase.

TABLE 9 Summary of gas phase reactor performances in “cobra’ (15flexible steel tube, 0.5 inch ID) packed-bed reactor. Total Acetol +Mass Acetol PG Glycerin PG Water Unknown balance Date Conditions Sample(wt %) (wt %) (wt %) (wt %) (wt %) (wt %) (wt %) G11 Direct Glycerin toPG with H2 Purge Oct. 22, 2005 Catalyst packing (size: Cobra 3 * 3 mm)760 g Hydrogen Purge Reactor temp: 240 C. Pressure: 27 in-Hg (vac) 155.13 13.91 10.88 69.04 16.4 96.32 Increase H2 Pressure: 19 in-Hg (vac)2 42.41 23.88 18.23 66.29 16.06 100.58 flow G12 Direct Glycerin to PGwith H2 Purge Oct. 24, 2005 Catalyst packing (size: Cobra 3 * 3 mm) 1160g Hydrogen Purge Reactor temp: 240 C. Pressure: 27 in-Hg (vac) 1 46.886.44 0 53.32 23.91 24.7 77.23 Increase H2 Pressure: 19 in-Hg (vac) 240.44 17.19 0 57.63 22.3 16.6 79.93 flow Increase H2 Pressure: 8 in-Hg(vac) 3 36.63 37.1 0 73.73 18.18 9.3 91.91 flow G13 Direct Glycerin toPG with H2 Purge Oct. 26, 2005 Catalyst packing (size: Cobra 3 * 3 mm)1160 g Hydrogen Purge Reactor temp: 230 C. Pressure: 27 in-Hg (vac) 158.56 5.58 0 64.14 64.14 excluding water Increase H2 Pressure: 19 in-Hg(vac) 2 44.94 25.07 0 70.01 70.01 excluding flow water Increase H2Pressure: 8 in-Hg (vac) 3 32.33 38.71 0 71.04 71.04 excluding flow waterG14 Direct Glycerin to PG with H2 Purge Oct. 25, 2005 Catalyst packing(size: Cobra 3 * 3 mm) 1160 g Hydrogen Purge Reactor temp: 220 C.Pressure: 27 in-Hg (vac) n n n n n n n Increase H2 Pressure: 19 in-Hg(vac) 1 44.58 29.9 0 74.48 19.71 94.19 flow Increase H2 Pressure: 8in-Hg (vac) 2 32.14 53.37 0 85.51 15.92 101.43 flow G15 Direct Glycerinto PG with H2 Purge Oct. 26, 2005 Catalyst packing (size: Cobra 3 * 3mm) 1160 g Hydrogen Purge Reactor temp: 210 C. Pressure: 27 in-Hg (vac)1 58.91 10.18 3.26 69.09 72.35 excluding water Increase H2 Pressure: 19in-Hg (vac) 2 44.2 28.92 4.3 73.12 77.42 excluding flow water IncreaseH2 Pressure: 8 in-Hg (vac) 3 32.56 41.33 8.39 73.89 82.28 excluding flowwater G16 Direct Glycerin to PG with H2 Purge Oct. 26, 2005 Catalystpacking (size: Cobra 3 * 3 mm) 1160 g Hydrogen Purge Reactor temp: 200C. Pressure: 27 in-Hg (vac) 1 56.13 3.53 2.11 59.66 61.77 excludingwater Increase H2 Pressure: 19 in-Hg (vac) 2 38.42 31.13 3.78 69.5573.33 excluding flow water Increase H2 Pressure: 8 in-Hg (vac) 3 28.5441.8 9.93 70.34 80.27 excluding flow water

As seen by the data G11 through G16, in every instance the increase inhydrogen pressure resulted in better closure of the mass balance (higherselectivity) and higher conversions from acetol to propylene glycol. At240° C., the selectivity to acetol/propylene glycol was lower than at220° C.—at this higher temperature more “junk” peaks were on the GCindicating that product was lost to undesirable side-products.Selectivity increased as the progressively as the temperature waslowered from 240° C. (G12) to 230° C. (G13) to 220° C. (G14). At 210° C.(G15) and 200° C. (G16), the conversion of glycerin was less than 100%.The optimal temperature is near 220° C.

The preferred operation of the evaporator is at a temperature near 230°C., and contact of glycerin with gases is such that glycerin attains apartial pressure of about 0.15 bars. A preferred stoichiometric additionof hydrogen feed 1036 will add an additional partial pressure of about0.15 bars or more of hydrogen.

It is possible to operate a two-reactor system such that acetol ispredominantly formed in the first reactor and acetol is predominantlyconverted to propylene glycol in the second reactor. The first reactorcould be a reactive distillation reactor operated at lower pressures orit could be a packed-bed reactor operated at lower pressures. There islittle advantage to using higher pressures in the first reactor if thegoal is to form acetol; the acetol is preferably produced from a processoperated at a pressure less than 5 bars. Higher pressures are preferredfor the hydrogenation of acetol to propylene glycol; for this reactionthe preferred pressure is greater than 25 bar and acetol may be presentas a liquid phase at these high pressures. When liquids are present,higher pressures are preferred, from at least 1 bar up to 500 bar orhigher. If liquid acetol is present in the reactor, the reactor ispreferably a slurry batch reactor, trickle bed reactor, and teabagreactor. The temperature for the hydrogenation is preferably less than220° C. A gas phase packed bed reactor may also be used for theconversion of acetol to propylene glycol.

The advantage of using a reactor where acetol is a liquid for convertingacetol to propylene glycol is that the liquid product can be selectivelyremoved (rather than hydrogen) from the reaction environment. This canreduce separation costs and the costs of recycling hydrogen. Anadditional advantage is that the high pressures and lower temperatures(less than 220° C.) can substantially overcome the equilibriumlimitations of the acetol-to-propylene glycol hydrogenation. Theadvantage of using a gas-phase packed-bed reactor is that lowerpressures are required. An additional advantage of using a gas-phasepacked-bed reactor is that by-product formation tends to be promoted inliquid phases—especially the formation of tar.

As an alternative to the two-reactor system, it is possible to formpropylene glycol in a single reactor by operating this reactor withsufficient hydrogen present. Tables 7 and 8 illustrate how 1 to 4 barsof hydrogen pressure are sufficient for high selectivity. The dataclearly extrapolates to favorable formation of acetol at lower hydrogenpressures (corresponding to lower partial pressures of hydrogen) andfavorable formation of propylene glycol at higher pressures(corresponding to high partial pressures of hydrogen).

Alternatively, it is possible to use a two-reactor system where thefirst reactor is operated to promote conversion of glycerolpredominantly to propylene glycol and the second reactor is operated atlower temperatures that create equilibrium favorable for furtherconversion of acetol to propylene glycol. Here, the first reactorpreferably has reaction temperatures greater than 200° C. and the secondreactor preferably has temperatures less than 220° C. Preferably theglycerol concentration entering the second reactor is less than half theglycerol concentration entering the first reaction step. Morepreferably, the first reaction step has reaction temperatures greaterthan 210° C. and the second reaction step has temperatures less than210° C.

For gas phase reactions operated at moderate pressures (less than 25bar), it is important to avoid formation of liquids. Liquids exhibitincreased rates of by-product formation. At a given pressure, there is aminimum temperature below which liquids form. This temperature is afunction of composition. The least volatile component is glycerol, andso, higher glycerol concentrations necessitate operating at highertemperatures, hence, in a two-reactor system the first reactor isoperated at a higher temperature because the first reactor has thehigher glycerol concentration.

The partial pressure of glycerol is the key parameter to be followed andcontrolled to avoid formation of liquids in the gas-phase packed-bedreactor. The partial pressure of glycerol must be kept below glycerol'sdew point partial pressure. Glycerol's dew point partial pressure isdefined as the partial pressure of glycerol below which glycerol doesnot form dew and at which a dew is formed that contains glycerol—thisdefinition implies that the glycerol-free ratios of all other componentsin the system remain constant and the temperature is constant as thepartial pressure of glycerol is increased until the a dew is formed.Being defined in this manner, glycerol's dew point partial pressure is astate property that is a function of temperature and the glycerol-freeconcentration of other components in the gas.

The preferred processes of this embodiment have a partial pressure ofglycerol less than glycerol's dew point partial pressure in the reactionmixture and greater than one fourth the dew point partial pressure inthe reaction mixture. More preferably, the partial pressure of glycerolis greater than half the dew point partial pressure in the reactionmixture. The reaction temperature is preferably below 230° C. because athigher temperatures, byproduct formation is favored. When thetemperature gets too low (such as 180° C.), glycerol's dew pointpressure becomes so low that reactor throughput can become too low to beeconomically viable. Also, at lower temperatures the dehydration ofglycerol becomes slower, further leading to increased reactor sizes.

The partial pressure of glycerol is approximately the mole fraction ofglycerol in the gas phase times the total pressure. Furthermore, forrelatively pure glycerol in a gas entering the reactor, the dew formedat glycerol's dew point partial pressure is relatively pure glycerol.Hence at the dew point, to a good approximation, y^(Dew) _(Glyc)P=P^(Sat) _(Glyc). where y^(Dew) _(Glyc) is the dew point mole fractionof glycerol in the gas phase, P is total pressure, and P^(Sat) _(Glyc)is the saturation pressure of pure glycerol at the temperature of thesystem. The dew point mole fraction of glycerol is the maximum that canbe held in the gas phase. Table 10 summarizes dew point mole fractionsat 1, 10, and 100 bar total pressure.

TABLE 10 Maximum mole fractions of glycerol in system at indicatedtemperature above which liquids will form in system. Vapor PressureMaximum y_(Glyc) Maximum y_(Glyc) Maximum y_(Glyc) Glycerol (Dew Point)(Dew Point) (Dew Point) T (° C.) (bar) @ 1 bar @ 10 bar @ 100 bar 1500.007 0.007 0.0007 0.00007 180 0.025 0.025 0.0025 0.00025 200 0.0580.058 0.0058 0.00058 220 0.126 0.126 0.0126 0.00126 250 0.345 0.3450.0345 0.00345

If the reactor feed of a gas phase reactor is comprised mostly ofhydrogen and glycerol, at higher pressures and lower temperatures,extremely large quantities of hydrogen need to be removed from theproduct. If this hydrogen is not to be lost, it needs to be recycled.This recycling can be very costly. Reasonably larger mole fractions athigher temperatures are problematic because by-products rapidly form attemperatures greater than about 230° C. A system analyses reveals thatpreferred conditions for gas phase glycerol reaction. Optimizationbalances high versus low temperature to achieve high reaction rates(higher temperatures) without high by-product formation (lowertemperature). Optimization also balances more-favorable equilibrium andreduced by-product formation (higher pressure) with maintainingreasonably-high concentrations of glycerol in the feed (lower pressure).For the gas phase reaction, the optimal conditions for glycerolconversion without by-product formation tends to be 210 to 230° C. and 2to 10 bars.

It is possible to operate gas phase reactors a very low concentrationsof glycerol with the remainder predominantly hydrogen; however, adisadvantage of this approach is that the hydrogen rapidly carries theglycerol through the packed-bed reactor leading to low catalystproductivity and high catalyst costs. FIG. 9 and Table 11 summarize theimpact of the molar ratio of hydrogen to glycerol on catalystproductivity (measured in grams of propylene glycol produced per hourper gram of catalyst in the packed-bed reactor). The maximum catalystproductivity is at a mole fraction of glycerol right below the dew pointmole fraction of glycerol. The previously indicated preferredconcentrations of glycerol expressed in terms of partial pressureindirectly specify the preferred molar ratios—the partial pressurespecification is independent of total system pressure. All data is from1 to 4 bars total pressure with most of the data at 2 bars pressure.

TABLE 11 Effect of H₂: Glycerol mole ratio on catalyst productivity forthe glycerol to propylene glycol reaction. Pressure Reactor of HydrogenGlycerin Product Catalyst Temperature Discharge flowrate flow rateH2:Glyc flow rate PG Productivity [g [° C.] [bar] [L/min] [g/hr][mol:mol] [g/hr] [wt %] PG/g cat] 220 1 16.7 500.0 8.83 473.88 50.840.37 220 1 22.1 90.0 70.98 70.20 47.29 0.05 220 2 36.2 226.2 42.26195.43 56.31 0.17 220 2 24.7 90.0 192.17 28.74 63.71 0.03 220 2 24.790.0 106.82 51.92 65.75 0.05 221 2 5.0 201.6 7.02 185.00 56.51 0.16 2202 5.0 127.2 12.14 99.96 51.20 0.08 221 2 5.0 127.2 12.40 97.68 49.380.07 220 4 29.9 226.2 26.72 258.84 69.47 0.28 220 4 22.1 500.0 13.70386.76 65.83 0.39 238 1 2.5 99.5 8.78 71.56 27.31 0.03 241 1 2.5 99.56.58 99.90 27.73 0.04 240 1 5.0 198.0 6.83 191.23 31.35 0.09 237 2 5.0198.0 9.91 125.00 35.53 0.07 236 2 5.0 198.0 9.79 126.80 40.56 0.08 2402 5.0 198.0 7.52 170.80 31.80 0.08 240 4 5.0 198.0 12.70 95.23 30.910.05 240 4 10.0 390.6 7.63 336.24 51.36 0.27

It is possible to operate the gas phase reaction at higher pressureswith high selectivity by making sure the partial pressure of glyceroldoes not exceed glycerol's dew point partial pressure. The advantage ofthis approach is to use high hydrogen pressures to push the equilibriumfrom acetol to propylene glycol and possibly reduce the need for someproduct purification unit operations and possibly eliminate the need torecycle acetol. Here, a process for converting glycerol to a product athigh selectivity to propylene glycol and low selectivity to ethyleneglycol preferably includes a reaction by: contacting a gas phasereaction mixture containing no liquid and containing a partial pressureof glycerol between 0.01 and 0.5 bars of glycerol and a total pressurebetween 25 and 500 bars with a heterogeneous catalyst at a temperaturebetween 150° C. and 280° C. for a reaction time interval between 0.01 to60 seconds. The solid catalyst preferably contains an element of thesubgroups from Group I, Group VI, and/or Group VIII of the PeriodicTable. Preferably, the partial pressure of glycerol is less thanglycerol's dew point partial pressure in the reaction mixture andgreater than one fourth the dew point partial pressure in the reactionmixture. More-preferably, the process of claim 72 where the partialpressure of glycerol is greater than half the dew point partial pressurein the reaction mixture.

A gas phase reactor can be used to produce acetol where pressures beyond1 bar little advantage—pressures lower than one bar can present a hazardif air leaks into the system. Here, a process for converting glycerol toa product at high selectivity to a mixture of acetol and propyleneglycol and low selectivity to ethylene glycol preferably includes areaction comprising: contacting a gas phase reaction mixture containingno liquid and containing a partial pressure of glycerol between 0.01 and0.5 bars of glycerol and a partial pressure of hydrogen between 0.01 and5 bars of hydrogen with a heterogeneous catalyst at a temperaturebetween 150° C. and 280° C. for a reaction time interval between 0.01 to60 seconds. The solid catalyst preferably contains an element of thesubgroups from Group I, Group VI, and/or Group VIII. Preferably, thepartial pressure of glycerol is less than glycerol's dew point partialpressure in the reaction mixture and greater than one fourth the dewpoint partial pressure in the reaction mixture. More-preferably, theprocess of claim 72 where the partial pressure of glycerol is greaterthan half the dew point partial pressure in the reaction mixture. Theslight amount of hydrogen in the system reduces by-product formation.

At 220° C. and 1 to 2 bars of pressure, the most preferred catalystloading in the reactor is about 4 grams of catalyst per gram ofpropylene glycol produced per hour (a 4:1 ratio). Ratios from 1:0.4 to1:0.05 are preferred. Higher catalyst loadings at the previouslyindicated preferred partial pressures of glycerol can lead to excessiveproduct loss to by-products.

Use of Base Neutralization in Finishing Process

Reaction by-products from the gas phase reaction include multiple estersof propylene glycol and ethylene glycol. Ethylene glycol is at timesobserved in concentrations less than 2%. One method of purifying thereaction products is to convert the esters back to propylene glycol andethylene glycol (along with corresponding acid) using a base (thereaction is effective at room temperature as well as temperatures higherthan room temperature). After this hydrolysis reaction, the glycols canbe evaporated/distilled from the salt formed from reaction of the basewith the ester's acid. The glycols can than be purified by distillation.

Here, the preferred product purifying process includes the steps of:adding a base to the said propylene glycol product to achieve a pHgreater than 8.0 and distilling the propylene glycol from the producthaving a pH greater than 8.0. The base is preferably selected from thegroup comprised of sodium hydroxide, potassium hydroxide, and calciumoxide.

Packed-Bed Reactor Design

Temperature control is most important to maximize reaction selectivityto acetol and/or propylene glycol. Reactors and methods known in the artcan be used effectively, including but not limited to fluidized bedreactors and packed bed reactors. The preferred reactor is a novelreactor for use with highly exothermic reactions comprised of an outershell containing U-tubes with an orientation such that the U-end of theU-Tubes is facing upward, said shell having an upper removable headwhere catalyst is loaded between shell and tubes from the top byremoving the upper head. An inert packing may be placed in the lowestportion of the space between the shell and U-Tubes at a depth between 2and 24 inches.

The design recommendation is to use a reactor sizing and loading of 4 lbof catalyst for each lb/hr of PG production. It is anticipated that atleast 2000 lbs of propylene glycol (PG) is expected to be produced perpound of catalyst. Table 12 summarizes the catalyst loadings based onanticipated capacity.

TABLE 12 Summary of catalyst loading recommendations. Basis CatalystComment  60 M lb PG/yr  30,000 lb cat Total loading {0.25 lb PG/[lb cat× hr]} 100 M lb PG/yr  50,000 lb cat Total loading ″  25,000 lb/reactorInitial Loading Per Reactor, 75% full ″  33,000 lb/reactor Loading PerReactor, 100% full Fallback 100,000 lb cat Catalyst loading if all 3reactors full Capacity

An initial total catalyst loading of 30,000 lb (assumes 0.25 lb PG/[lbcat×hr])) (4 lb catalyst per lb/hr of PG production) may be appropriateif the initial capacity is targeted at 60 million lb per year.

One design uses three reactors arranged to operate in series with eachreactor capable of holding 33,000 lb of catalyst when full and with theintent of only operating two in series at any time. Initially, eachreactor would be 75% full and catalyst would be added as the initialloading deactivates—catalyst will be added until the reactor is full ofcatalyst. When catalyst activity is too low to meet production capacity,the first reactor in the series will be removed from operation and thereserve reactor will be placed as the second reactor in series with aninitial catalyst loading of 50-75%. Inert material should be used in thefirst reactor for initial startup. After this startup, no inert materialshould be required in any of the reactors because the catalyst will besufficiently deactivated in the second reactor by the time it is placedfirst in the series. This is with the exception that inert packing isalways to be placed in the bottom 6 inches of the reactor as it is inthis location were dead spots might occur relative to flow.

The design recommendation on the reactors is for use of tube-and-shellheat exchanger reactors that have U-tube internals. The reactors shouldbe configured with the U end of the tube upward and a flanged/removablehead on this upward U-tube side. Catalyst is to be loaded from the topwith precautions taken to assure catalyst level is even through out allthe spaces between U-tubes. More catalyst can be added to the reactor(added to top of catalyst bed) as the reaction proceeds. An upward flowis recommended. The packing is on the shell side (not tube side).

The desired catalyst loading to provide an ultimate capacity of 100million lb/yr of PG production is summarized in Table 13 and can be metwith 3, 6, or 9 reactors at the sizes and configurations indicated.

TABLE 13 Summary of reactor configuration options to meet catalystcapacity. Assumes ¾ inch tubes with wall-to-wall spacing of 1 inch intriangular configuration. Length of Shell Diameter of Shell # Reactors(ft) (ft) Description 3 12′ 7.5 3 in series 3 16′ 6.5 3 in series 6 12′5.3 Two sets of 3 in series 6 16′ 4.6 Two sets of 3 in series 9 12′ 4.3Three sets of 3 in series 9 16′ 3.8 Three sets of 3 in series

Using these specifications for reactors and a design intended to use 2of 3 reactors in series, the design productivity ratio is 0.25 [lb PG/lbcat-hr]), but the system is capable of producing at “worst case” ratesof 0.125 [lb PG/lb cat-hr]). This should be an ample safety factor toguarantee the design capacity.

Reactor Sizing

The catalyst density is 111 lb/ft³ (some error exists due to the non-drynature of the catalysts, all masses assume that free liquids are removedbut catalyst surfaces are moist). This translates to 33,300 lb/111lb/ft³, or 300 cubic feet of void volume per reactor. In tube-and-shellheat exchanger design with ¾″ OD tubes on a 1″ (wall-to-wall) triangularspacing, the void volume is estimated to be 83%. Assuming a slightpacking inefficiency, a value of 75% void volume (shell volume availablefor packing, not occupied by tubes) is warranted. Thus, the volume ofthe shell (not including heads) should be 400 cubic feet.

An area of 400 cubic feed divided by 12 ft length (length of heatexchanger) translates to a void cross sectional areas of 33.3 ft²(400/12) or shell cross sectional areas (75% void volume) of 44.4 ft².This translates to a shell of 7.5 ft ID. The area in inches is 6400 in².Under the assumption of one tube for each 1.33 in², each heat exchangerwould have 4828 tubes of ¾″ OD (or 2414 U-tubes). Results from thesecalculations are summarized in Table 12.

Data on Catalyst Loading

FIG. 10 shows results from such a reactor as modeled by COBRA software.For the multi-tube reactor, the catalyst productivity is approximately625 grams of PG/hr divided by 3.6 kg of catalyst (or 0.174 [lb PG/lbcat-hr]). Comparing COBRA to multi-tube reactor performances, it isconcluded that the performance in the current multi-tube reactor islimited by the evaporator and not the amount of catalyst. In addition,acetol recycle is expected to increase catalyst productivity by up to50%. For these reasons, the recommended catalyst loading is 0.25 [lbPG/lb cat-hr] (4 lb of catalyst per 1 [lb/hr] PG production) rather than0.174 [lb PG/lb cat-hr].

The mass of catalyst loading is calculated by taking the hourlyproduction rate of PG and multiplying this number times 4 hrs. The hoursof operation in a year are assumed at 360×24=8640 hours/yr or a plantcapacity of 11,574 lb/hr. This translates to a catalyst loading of46,300 lb for a 100 million lb/yr facility (11,574×4).

One type of design is for three reactors in series with each reactorcapable of holding 33,000 lb of catalyst when full. The intent is toonly operate two reactors in series at any time. Initially, each reactorwould be 75% full and catalyst would be added as the initial loadingdeactivates. Table 13 summarizes six different ways to achieve thiscapacity.

Comments on Shell-Side Loading

A shell side loading of the reactor is possible because access from thetop (U-side) allows easy loading of catalyst in the reactor allows inertpacking to be strategically placed in the reactor. As shown in FIG. 11A,a shell-side reactor 1100 has a feed or inlet port 1102, that within theconfines of reactor 1100 is in fluidic communication with exit or outletport 1104. the reactor 1100 is used to perform the reactions describeabove. As shown in FIG. 11B, the reactor 1100 has an outer shell 1105that may be filled with one or more U-tubes i106 for use as heatexchangers within a jacket 1108 A top 1110 (see FIG. 11A) may be removedto permit top access to the jacket and to the reactor interior. core. Asshow in FIG. 11 C. the reactor interior core 1112 may be filled withmaterial including a bed of insert packing 1114 beneath a catalyst 1116,as described above. The advantages of this configuration over atube-side loading are as follows:

-   -   Up to a 5× reduction in reactor size is attained.    -   Potential problems of by-passing and hot-spot development that        can occur with tube reactors is all but eliminated.    -   The pressure drop is considerably less.    -   Filling catalyst in the tubes is considerably easier. Basically        2-6 shells are loaded with catalyst as opposed to 50,000 tubes        (5× volume).

EXAMPLE 4 Pilot Scale Reactor Validation

Pilot scale reactor validation of the “shell-side” approach to packingthe catalyst was successful and is considered fully scalable. This pilotreactor had a 2″ ID with three 0.5″ OD tubes in which oil was circulatedto cool the reactor. Various operational aspects validated in the pilotproduction run included the following:

-   -   Temperatures can be controlled (25% inert loading was used, may        or may not be necessary) with distances up to 1 inch from the        wall of the tube to the furthest catalyst particles. Conversions        were good and product profiles consistent with previous runs.    -   The system can operate with adequately low pressure drops.

One problem that developed was some (˜5%) glycerol vapors wereby-passing the catalysts and were in the product stream at conditionsthat previously exhibited less than 0.2% glycerol in the product. Theby-passing was believed due to the low depth of the reactor bedpacking—a packing of 3.75 feet of catalyst (5 feet of catalyst plusinert) as compared to 16 or 24 feet of catalyst (axial length ofpacking) in the cobra reactors. In a commercial facility, the depth maybe 18 to 24 feet, and so, this bypass issue should be resolved.

FIG. 12 shows an axial and radial cross section of the pilot reactor1200. The heat transfer in this reactor is over the same dimensions asheat exchange in the reactors specified above The reactor 1200 is fullyscalable by simply increasing the shell diameter while keeping thespacing of the oil-filled circulating heat exchange tubes 1204, 1206,1208 the same. The shell surface 1210 was insulated against heat loss.Reactor feed inlet 1212 and reactor feed outlet 1214 are in fluidiccommunication through a catalyst-packed chamber 1214.

The following three pilot reactors have been demonstrated in thelaboratory with the first two having tube loadings and the last having ashell loading:

-   -   16′ pilot using ¾″ OD tube: 1500 g catalyst loading.    -   10′ pilot using 1.0″ OD tube: 2280 g catalyst loading.    -   5′ pilot using 2.0″ ID shell and 3 tubes: 3600 g catalyst        loading

Even at the pilot scale, the shell-side loading was much easier to workwith and the reactor had a much lower pressure drop.

TABLE 14 Pilot Scale Results Shell-and-Tube reactor Shell-and-Tubereactor Tube-cooled reactor #1 (¾″ OD; #2 (1″ OD; (2″ ID; 16 ft. length)10 ft. length) 5 ft. length) Glycerol flow Unconverted UnconvertedUnconverted rate (L/hr) Color Glycerol (%) Color Glycerol (%) ColorGlycerol (%) 0.3 clear 0 — — — — 0.5 clear 0 — — — — 0.75 clear 0 clear0 clear (not include) 0.8 Very 1 clear 0 clear ~3 slightly yellow 1Light yellow 5 Very (not include) clear ~8 slightly yellow 1.3 Yellow 8Light yellow (not include) clear ~14 

A preferred pressure drop from compressor exit to compressor entrance is14.7 psi. The pilot scale system has operated for several days at apressure drop of 7.5 psi. This validates the capability to operate atthe needed low pressure drop.

In the pilot facility, about 5 psi pressure drop is across theevaporator, about 0.5 psi pressure drop is across the 5 feet of reactorpacking, and about 2 psi pressure drop is from the condenser andentrance/exit effects. Increasing the bed depth to 20′ should result in2 psi. A further doubling of space velocity through the packing shouldincrease the pressure drop to 4 psi. A 4 psi pressure drop throughpacking is consistent with targeted commercial reactor operation.

The 5 psi pressure drop in the evaporator is consistent with a “long”travel path of at least 100 feet of evaporator tubing. The targetedevaporator pressure drop for the commercial system is 3 psi. Thetargeted pressure drop for all other aspects (condenser, flash vessels,and entrance effects) is 3 psi.

Composition of Antifreeze Product from Glycerol of Biodiesel Facility

Biodiesel is one type of product that can be produced from a fatty acidglycerin refinery. After a conventional biodiesel methanolysis reaction,the methoxylation catalyst is preferably removed by filtration from aslurry reaction system. Other methods, such as centrifugation orprecipitation, may be used to remove soluble catalysts from the glycerolby-product of the biodiesel methanolysis reaction process. Theseprocesses are compatible with either batch or continuous operation.Methods known in the art may be used to convert the batch processprocedures (described herein) to flow process procedures. Hydrogenationof the glycerol is performed to prepare a glycerol byproduct thatpreferably contains, on a water-free basis, from 0.5% to 60% glycerol,and 20% to 85% propylene glycol. More preferably, the glycerol byproductcontains on a water-free basis from 10% to 35% glycerol, and 40% to 75%propylene glycol. Also, as the preferred antifreeze of this invention isprepared from the crude natural glycerol byproduct of the C₁ to C₄ alkylalcohol alcoholysis of a glyceride, the more preferable product alsocontains 0.2% to 10% C1 to C4 alkyl alcohol and 0 to 5% salt of theneutralized alcoholysis catalyst (more preferably 0.2 to 5% salt).

The glycerol conversion reactions have been observed to form a residueby-product. When this residue is soluble in the antifreeze product, thepreferred application is to add it to the antifreeze product. Theantifreeze may contain 1% to 15% of this residue by-product.

While the antifreeze products of this invention are commonly referred toas antifreeze, these same mixtures or variations thereof may be used asdeicing fluids and anti-icing fluids.

When the reaction is run without hydrogen, acetol will form. Thismixture can then subsequently (or in parallel) react in a packed-bedflow reactor in the presence hydrogen to be converted to propyleneglycol. This process has the advantage that the larger reactor does notcontain pressurized hydrogen.

The processes and procedures described in this text are generallyapplicable to refined glycerol as well as crude glycerol.

The catalyst used for most of the process development was a Sud-Chemiepowder catalyst at 30 m²/g surface area, 45% CuO, 47% Cr₂O₃, 3.5% MnO₂and 2.7% BaO. Also used was a Sud-Chemie tablet catalyst at 49% CuO, 35%Cr₂O₃, 10% SiO₂ and 6% BaO. Also used was a Sud-Chemie powder catalystat 54% CuO and 45%.

EXAMPLE 5

Processing of Biodiesel Byproduct

Crude glycerol obtained as a by product of the biodiesel industry wasused instead of refined glycerol. Biodiesel is produced usingalcoholysis of bio-renewable fats and oils. The composition of feedstock104 used in this example had an approximate composition as follows:glycerol (57%), methyl alcohol (23%), and other materials (soaps,residual salts, water) (20%). The above feedstock was reacted in thepresence of hydrogen and catalyst to form a mixture containing propyleneglycol. The reaction proceeded using 10 grams of the crude feedstock, 5%by weight of catalyst, and a hydrogen overpressure of 1400 kPa. Thefollowing Table 15 presents compositions after reacting in a closedreactor (with topping off of hydrogen) for 24 hours at a temperature of200° C. The copper chromium catalyst used in this Example was reduced inpresence of hydrogen at a temperature of 300° C. for 4 hours prior tothe reaction.

TABLE 15 Summary of catalyst performances based on 10 grams of crudeglycerol. Initial Loading Best Final Product (g) Possible (g) (g) CrudeGlycerol 5.7 0 0.8 Acetol 0 0 0 Propylene glycol 0 4.6 3.1 Water 1 2.12.6Reactive-Separation to Prepare Acetol and Other Alcohols

As an alternative to reacting to form propylene glycol by use of theprocess equipment 100 shown in FIG. 1, FIG. 13 shows a modified versionof the process equipment that has been previously described. Processequipment 1300 is useful for forming acetol or other alcohols havingboiling points less than about 200° C. Dehydration is the preferredreaction method, but cracking reactions may be used with feed stockscontaining sugars or polysaccharides having carbon numbers greater than3.

In general, the process equipment 1300 is used for converting athree-carbon or greater sugar or polysaccharide to an alcoholdehydration product having a boiling point less than about 200° C. Byway of example, a sugar or polysaccharide-containing feedstock with lessthan 50% by weight water is combined with a catalyst that is capable ofdehydrating glycerol to form a reaction mixture. The reaction mixture isheated to a temperature ranging from 170° to 270° C. over a reactiontime interval ranging from 0.5 to 24 hours at a pressure ranging from0.2 to 25 bar.

One method of converting glycerol to acetol is in a semi-batch slurryreaction where glycerol is continuously added and acetol is removed as avapor. In this method, residence time may be less than 0.5 hours,depending upon catalyst loading. Also, pressures as low as 0.02 bar willsuccessfully pull off the acetol as a vapor that can then be condensedto form a liquid product.

The preferred reaction conditions for conversion of glycerol to formacetol include a process temperature ranging from 170° C. to 270° C.,and this is more preferably from 180° C. to 240° C. The preferredreaction time exists within a range from 0.5 to 24 hours. Heterogeneouscatalysts that are known to be effective for dehydration may be used,such as nickel, copper, zinc, copper chromium, activated alumina andothers known in the art. The preferred reaction pressure exists within arange from 0.2 to 25 bar, and this is more preferably from The 0.5 to 3bar. The feedstock may contain from 0% to 50% and more preferably 0 to15% water by weight.

By these instrumentalities, glycerol may be dehydrated to acetol.Selective formation of acetol is documented for the copper-chromiumcatalyst by Examples 5 through 7 below. The same reaction conditionswith different catalyst are effective for forming other alcohol productswhere the products have fewer alcohol functional groups than do thereagents. Fractional isolation of intermediates throughreactive-distillation is particularly effective to increase yields andthe embodiments is inclusive of processes to produce a range of productsincluding but not limited to 1,3 propanediol and acrolein.

FIG. 13 shows process equipment 1300 for the selective conversion ofglycol to acetol. In FIG. 13, identical numbering is used for the samecomponents that have been previously described with respect to FIG. 1.The reactor separator 102 as shown in FIG. 13 functions as a dehydrationreactor. The polyhydric feedstock 104 and catalyst 124 enterreactor-separator 102 for a reaction that is limited to the dehydrationstep 204 of FIG. 2 by the absence of hydrogen, and in consequence thehydrogenation step 208 does not occur at this time. The dominantreaction product is acetol 202. Volatile fractions including acetolvapor exit the reactor-separator 102 through an overhead intermediateline 1302 and liquefy in condenser 112. A follow-on reactor 1304functions as a hydrogenolysis reactor that accepts acetol and otherliquids from condenser 112, and contacts the acetol with hydrogen toform propylene glycol as product 118. The catalyst 1306 may be the sameas or different from catalyst 126. The condenser 112 preferably operatesat a temperature ranging from 25° C. to 150° C. and this ismore-preferably from 25° C. to 60° C. It will be appreciated that thecondenser 112 may be eliminated or positioned downstream of thefollow-on reactor 1304 if the follow-on reactor 1304 operates as a vaporphase reactor.

When the process equipment 1300 is operating in mode of producingpropylene glycol product 118, a hydrogen recycle loop 1312 recyclesexcess hydrogen from the follow-on reactor 1304. This step preferablyrecycles unused hydrogen from the condenser back to the subsequent stepreaction mixture. The reaction time of this subsequent step reactionexists within a range from 0.2 to 24 hours and more-preferably existswithin a range from 1 to 6 hours.

The acetol that is delivered through intermediate line 1302 to condenser112 is optimally diverted through three way valve 1308 to provide anacetol product 1310.

EXAMPLE 5 Stepwise Production of Acetol then Propylene Glycol

Glycerol was reacted in the presence of copper chromium catalyst in twosteps to form a mixture containing propylene glycol. In Step 1,relatively pure acetol was isolated from glycerol in absence of hydrogenat a reaction pressure of 98 kPa (vac). In Step 2, the acetol from Step1 was further reacted in presence of hydrogen to propylene glycol at1400 kPa hydrogen over pressure using similar catalyst that is used forthe formation of acetol. The catalyst used in the step 1 of this Exampleis used in the condition in which they arrived and the catalyst used inthe Step 2 was reduced in presence of hydrogen at a temperature of 300°C. for 4 hours prior to the reaction.

The following table presents composition of the final product in Step 1and Step 2

TABLE 16 Example reaction conditions for converting glycerol topropylene glycol. Initial Loading Best (g) Possible (g) Final Product(g) Step 1: Formation and isolation of acetol intermediate from glycerolusing copper-chromite catalyst. Catalyst—5% unreduced powder Cu/Cr,Reaction time—1.5 hr at 220° C. and 3 hr at 240° C., ReactionPressure—98 kPa (vac). Glycerol 36.8 0 3.6 Acetol 0 29.6 23.7 Propyleneglycol 0 0 1.7 Water 0 7.2 6.9 Step 2: Formation of propylene glycolfrom acetol intermediate from Step 1 using same catalyst. Catalyst—5%reduced powder Cu/Cr, Reaction time—12 hr, Reaction Temperature—190° C.,Reaction Pressure—1400 kPa. Glycerol 0 0 0 Acetol 4.5 0 0 Propyleneglycol 0 4.6 4.3

EXAMPLE 5 Batch Versus Semi Batch Processing

Glycerol was reacted in presence of copper chromium catalyst to formacetol by each of two process modes: batch and semi batch. Relativelypure acetol was isolated from glycerol in absence of hydrogen at areaction pressure of 98 kPa (vac). In this reaction 92 grams of glycerolwould form a maximum of 74 grams acetol at the theoretical maximum 100%yield. Either process mode produced a residue. When dried, the residuewas a dark solid coated on the catalyst that was not soluble in water.

In semi-batch operation, the reactor was provisioned with catalyst andglycerol was fed into the reactor at a uniform rate over a period ofabout 1.25 hours. In batch operation, all of the glycerol and catalystwas loaded into the reactor at the start of the reaction. The followingresults show the semi-batch reactive-distillation has higher yields andselectivities than batch. The higher catalyst loading provided higheryields and selectivities. It was observed that the catalyst activitydecreased with reaction time and the amount of residue increased withreaction time.

The copper chromium catalysts used in this Illustrative Example wereused in the condition in which they arrived. Process runs were madeusing the conditions described in Table 17.

TABLE 17 Comparison of Semi-Batch (Continuous) Reactive-distillation andBatch Reactive-distillation. Formation and isolation of acetolintermediate from glycerol using copper-chromite catalyst. Catalyst—5%unreduced copper chromium powder Reaction conditions: ReactionPressure—98 kPa (vac) Reaction temperature—240° C. Reaction completetime—2 hr Glycerol feed rate—33.33 g/hr for Semi-Batch Reactions Thefollowing three reactions were conducted: RXN 8.1—Semi-Batch reaction at5% catalyst loading RXN 8.2—Semi-Batch reaction at 2.5% catalyst loadingRXN 8.3—Batch reaction 5% catalyst loading

Table 18 provides reaction details of reaction conditions RXN 8.1 ofTable 17. Initial loading of glycerol was 54.29 grams; glycerol indistillate was 4.91 grams; residue was 3.80 grams, and the amount ofglycerol reacted was 49.38 grams.

TABLE 18 Mass balance details on RXN 8.1. Catalyst loading was 5%.Reacted Glycerol (g) Best possible (g) Distillate (g) Glycerol 49.38 03.64 Acetol 0 39.71 35.99 Propylene 0 0 1.65 glycol Water 0 9.66 5.79

Table 19 provides details of reaction 8.2: Initial loading of glycerolwas 52.8 grams; glycerol in distillate was 3.85 grams; residue was 4.91grams; and the amount of glycerol reacted was 48.95 grams.

TABLE 19 Mass balance details on RXN 8.2. Catalyst loading was 2.5%.Reacted Glycerol (g) Best possible (g) Distillate (g) Glycerol 48.95 03.85 Acetol 0 39.37 33.51 Propylene glycol 0 0 1.63 water 0 9.58 6.24

Table 20 provides details of reaction 8.2: Initial loading of glycerolwas 42.48 grams; glycerol in distillate was 3.64 grams; residue was5.68; and the amount of glycerol reacted was 33.16 grams.

TABLE 20 Mass balance details on RXN 8.3. Catalyst loading was 5%.Reacted Glycerol (g) Best possible (g) Distillate (g) Glycerol 36.80 03.64 Acetol 0 29.60 23.73 Propylene glycol 0 0 1.67 water 0 7.2 6.99

As reported in the following examples, various studies were performed toassess the ability to control the residue problem.

EXAMPLE 6 Control of Residue by Water Content of Feedstock

Glycerol was reacted in presence of copper chromium catalyst to formacetol at conditions similar to Illustrative Example 5 with 2.5%catalyst loading and in a semi-batch reactor method. Water was added tothe glycerol to evaluate if water would decrease the accumulation of thewater-insoluble residue. Table 21 summarizes the conversion results.These data illustrate that a small amount of water reduces the tendencyfor residue to form. The copper chromium catalyst used in thisIllustrative Example was used in the condition in which they arrived.

TABLE 21 Impact of water on residue formation. Catalyst—2.5% unreducedpowder Cu/Cr Reaction Pressure-98 kPa (vac) Reaction temperature-240° C.Reaction complete time-2 hr Glycerol feed rate-33.33 g/hr Water InitialGlycerol in Best Possible of Acetol in Distillate Residue:Initial (wt %)Glycerol (g) Distillate (g) Acetol (g) (g) Residue (g) Conversion (%)Glycerol Ratio  0% 52.8 3.85 39.37 33.51 4.91 92.71% 9.30%  5% 53.264.93 38.87 35.23 3.47 90.74% 7.02% 10% 56.25 8.55 38.36 34.48 3.4584.80% 6.13% 20% 55.52 9.67 36.87 33.13 2.95 82.58% 5.31%

EXAMPLE 7 Control of Residue by Catalyst Loading

Glycerol was reacted in presence of copper chromium catalyst to formacetol in a semi-batch reactor method. The impact of lowering catalystloadings was evaluated to determine the impact of catalyst loading onacetol yield and residue formation. Table 22 summarizes the conversionresults. These data illustrate that the formation of residue may beautocatalytic—it increases more than linearly with increasing throughputof glycerol over the catalyst. Also, the selectivity decreases withincreasing throughput of glycerol over a fixed catalyst loading in thereactor.

The copper chromium catalyst used in this Illustrative Example was usedin the condition in which they arrived.

TABLE 22 Impact of catalyst to glycerol throughput ratio on residueformation. Catalyst—1.25 g unreduced powder Cu/Cr Reaction Pressure-98kPa (vac) Reaction temperature-240° C. Glycerol feed rate-33.33 g/hrTotal feed of Acetol Residue:Reacted- Reaction Catalyst % Glycerol (g)Residue (g) Conversion (%) Selectivity Glycerol Ratio 1   5% 27.15 1.990.96% 90.62% 7.70% 2 2.50% 52.80 4.91 92.71% 85.11% 10.03% 3 1.67%77.22 7.54 90.44% 76.94% 10.76% 4 1.25% 105.68 11.7 89.23% 73.50% 12.11%5 0.83% 151.69 17.18 86.87% 59.76% 13.03%

EXAMPLE 8 Regeneration of Catalyst

This example illustrates the stability of the copper chromium catalystfor the formation of propyelene glycol from acetol. The following werethe approximate initial conditions: 4.5 grams of acetol, 2 wt % ofcatalyst, and a hydrogen overpressure of 1400 kPa. The following tablepresents compositions after reacting in a closed reactor (with toppingoff of hydrogen) for 4 hours at a reaction temperature of 185° C. Thecopper chromium catalyst was reduced in presence of hydrogen at atemperature of 300° C. for 4 hours prior to the reaction. The catalystafter each run was filtered from the reaction products, washed withmethanol and then dried in a furnace at temperature of 80° C. Thisregenerated catalyst was reused in the subsequent reactions. Similarregeneration procedure is repeated 10 times and the results aresummarized in Table 23. These data illustrate the ability to reusecatalyst for the hydrogenation of acetol.

The copper chromium catalyst used in this Illustrative Example wasreduced in presence of hydrogen at a temperature of 300° C. for 4 hoursprior to the reaction.

TABLE 23 Summary of catalyst performances based on 4.5 grams of acetol.Acetol (g) Propylene glycol (g) Lactaldehyde (g) initial 4.5 0 0 Run 10.5 3.62 0.51 Run 2 0.29 3.85 0.56 Run 3 0.19 4.19 0.53 Run 4 0.07 4.410.47 Run 5 0.05 4.42 0.49 Run 6 0.05 4.39 0.51 Run 7 0 4.41 0.36 Run 80.24 4.2 0.42 Run 9 0.27 4.2 0.43 Run 10 0.21 4.11 0.4

EXAMPLE 9 Ability to Reuse Catalyst of Acetol-Forming Reaction

This example illustrates that a powder catalysts may be treated orreactivated by hydrogen treatment, but also that one powder catalystthat contains 54% CuO and 45% Cr₂O₃ has better reuse properties thandoes another powder catalyst at 30 m²/g surface area, 45% CuO, 47%Cr₂O₃, 3.5% MnO₂ and 2.7% BaO. For the powder catalyst at 54% CuO and45% Cr₂O₃, the data of Table 24 demonstrate that residue formation rateis similar to that of the powder catalyst at 30 m²/g surface area, 45%CuO, 47% Cr₂O₃, 3.5% MnO₂ and 2.7% BaO (Table 23). The data of Table 25demonstrate the 54% CuO and 45% Cr₂O₃ catalyst can be used repeatedly(at laboratory scale, 1-3% of the catalyst was not recovered fromreaction to reaction). The data of Table 26 demonstrate that reuse ismore difficult with the 45% CuO, 47% Cr₂O₃, 3.5% MnO₂ and 2.7% BaOCatalyst.

TABLE 24 Impact of catalyst to glycerol throughput ratio on residueformation. The catalyst in this table is a powder catalyst at 54% CuOand 45% Cr₂O₃. This compares to the catalyst of Table 13 which is apowder catalyst at 30 m²/g surface area, 45% CuO, 47% Cr₂O₃, 3.5% MnO₂and 2.7% BaO. Reactions were semi-batch. Catalyst—1.25 g unreducedCu/Cr, powder catalyst at 54% CuO and 45% Cr2O3. Pressure-98 kPa (vac);Temperature-240° C.; Glycerol feed rate-33.33 g/hr Total feed Acetol[Residue]:[Reacted- of Glycerol Selectivity Glycerol] Reaction Catalyst(%) (g) Residue (g) Conversion (%) (%) Ratio 1   5% 26.35 1.95 89.82%87.05% 8.36% 2 2.50% 53.38 5.41 91.05% 82.01% 11.13% 3 1.25% 102.9812.36 89.07% 78.86% 13.47%

TABLE 25 Impact of reuse on powder catalyst at 54% CuO and 45%. Catalystis loaded at 5% and is unreduced. Catalyst—2.5 g unreduced Cu/Cr, powdercatalyst at 54% CuO and 45% Cr₂O₃. Pressure-98 kPa (vac);Temperature-240° C.; Glycerol feed rate-33.33 g/hr Con- Acetol Totalfeed of Residue version Selectivity Residue:Initial- Glycerol (g) (g)(%) (%) Glycerol Ratio Fresh 52.77 3.96 89.82% 87.05% 7.51% Reused 152.16 4.11 91.28% 88.52% 7.88% Reused 2 51.72 3.89 91.74% 88.56% 7.53%Reused 3 Catalysts still could be recovered

TABLE 26 Impact of reuse on powder catalyst at 30 m2/g surface area, 45%CuO, 47% Cr2O3, 3.5% MnO2 and 2.7% BaO. Catalyst—2.5 g unreduced powderCu/Cr Pressure-98 kPa (vac); Temperature-240° C.; Glycerol feedrate-33.33 g/hr Total feed Con- Acetol of Glycerol Residue versionSelectivity Residue:Initial- (g) (g) (%) (%) Glycerol Ratio Fresh 54.293.80 90.95% 90.62% 7.01% Reused 1 53.13 3.99 88.92% 88.80% 7.51% Reused2 Catalyst could not be recovered-residue was totally solidified

The two catalysts at initial condition performed about the same for theacetol forming reaction; however, the 45% CuO, 47% Cr₂O₃, 3.5% MnO₂ and2.7% BaO catalyst at a loading of lesser than 5% formed a different typeof residue that was more resistant to catalyst recovery. For bothcatalysts, it was generally observed that as reactions proceeded, thereaction rates tended to reduce. At the end of the semi-batch reaction adigestion of the mixture was induced by stopping the feed and allowingthe reaction to proceed for about 30 min to an hour—during thisdigestion the volume of the reaction mixture decreased and the residuebecame more apparent.

For the 54% CuO and 45% Cr₂O₃ catalyst, the residue tends to be stable.This residue takes a solid form in room temperature and a slurry form atthe reaction temperature during the long period of reaction time. Amethanol wash readily removed the residue, allowing the catalyst to bereused multiple times. The solid was soft and tacky in nature andreadily dissolved in methanol to form slurry. The catalyst was washedwith methanol until the wash was clear and then the catalyst was driedin a furnace at 80° C. to remove the methanol. The physical appearanceof this catalyst after washing was similar to that of the new catalyst.

In the case of 45% CuO, 47% Cr₂O₃, 3.5% MnO₂ and 2.7% BaO catalyst theresidue was, however, different. In the case of 5% catalyst loading,residue started foaming on the catalyst at 30 min after total glycerinwas fed, i.e., 30 minutes into the reaction. Once foaming started, amethanol wash was not effective for removing the residue from thecatalyst. If the reaction was stopped prior to commencement of foaming,the methanol was effective in removing the residue from the catalyst.When catalyst loading less than 2.5%, the residue started foaming whilethe glycerin was still being fed to the reactor, and the catalyst couldnot be recovered at end of the reaction. The 54% CuO and 45% Cr2O3catalyst produced a residue that is a solid at room temperature.

These trends in reuse of catalyst are applicable to conditions forconversion of glycerin to acetol as well as the “single-pot” conversionof glycerin to propylene glycol.

EXAMPLE 10 Lactaldehyde Mechanism

Acetol was hydrogenated in presence of copper chromium catalyst to forma mixture containing propylene glycol. The following were theapproximate initial conditions: 10 grams of acetol, 2 wt % of catalyst,and a hydrogen overpressure of 1400 kPa. The following table presentscompositions after reacting in a closed reactor (with topping off ofhydrogen) for 4 hours at a reaction temperature of 190° C. Table 27shows the effect of reaction temperature on the formation of propyleneglycol from acetol. The data illustrate that good conversions areattainable at 190° C. The data illustrate that the co-product (likelyundesirable) of lactaldehyde is produced at lower selectivities attemperatures greater than 150° C. Optimal temperatures appear to be 190°C. or higher. The copper chromium catalyst used in this IllustrativeExample was reduced in presence of hydrogen at a temperature of 300° C.for 4 hours prior to the reaction.

TABLE 27 Summary of catalyst performances based on 9 grams of acetol.The pressure is 1400 kPa with a 5% catalyst loading. Temperature CAcetol (g) Propylene glycol (g) Lactaldehyde (g) Unreacted 10 0 0  508.25 1.86 0.13 100 5.74 3.93 0.47 150 3.10 4.31 2.82 180 1.64 7.90 0.89190 0.56 9.17 0.58

Table 28 shows the effect of initial water content in the reactants onthe formation of propylene glycol from acetol. The data illustrate thatwater can improve yields to propylene glycol. Selectivity to propyleneglycol decreases as the reaction goes beyond 10-12 hrs.

TABLE 28 Summary of catalyst performances based on different initialloadings of water. The reaction temperature is 190° C., at a pressure of1400 kPa, a 5% catalyst loading and a reaction time of 24 hours. Thetotal loading of water with acetol is 10 grams. Water (% wt) Acetol (g)Propylene glycol (g) Lactaldehyde (g) 10 0.47 7.65 0 20 0.22 5.98 0.7 500.22 4.35 0.27

Table 29 shows the effect of initial catalyst concentration on theformation of propylene glycol from acetol. The data illustrate that thehighest yields are attained at the higher catalyst loadings.

TABLE 29 Summary of catalyst performances based on 4.5 grams of acetol.The reaction temperature is 190° C., at a pressure of 1400 kPa, and noadded water. Catalyst Concentration Reaction Propylene Lactaldehyde (wt%) Time (h) Acetol (g) glycol (g) (g) Initial — 4.5 0 0 5% 4 0.29 4.460.22 2% 4 0.14 4.27 0.2 1% 4 1.32 3.45 0.29 0.5%   4 1.56 3.14 0.32 1% 60.58 3.78 0.25 0.5%   6 1.27 3.29 0.33Reactive-Separation with Gas Stripping

The use of the reactor-separator 102 is very effective for convertingglycerol to acetol as illustrated by the foregoing Examples. Theseexamples illustrate, for example, the effective use of water andcatalyst loading to reduce formation of residue. Two disadvantages ofthe reactions were the formation of residual and operation at smallamounts of vacuum.

The most preferred approach overcomes the vacuum operation by using agas to strip the acetol from solution. Thus, the process equipmentoperates at a more optimal pressure, such as slightly over atmosphericpressure, to make advantageous use of fugacity (partial pressures) forthe selective removal of vapor from the reaction mixture. The strippergases may be inert gases such as nitrogen to strip out the acetol. Steammay also be used to strip out the acetol. The most-preferred strippinggas is hydrogen.

Use of hydrogen at pressures slightly above atmospheric pressure stripsthe acetol and/or propylene glycol from solution as they are formed. Inaddition, the preferred hydrogen stripper gas keeps the catalyst reducedand provides reaction paths that prevent residue formation and/or reactwith the residue to form smaller molecules that also strip fromsolution. The reactions that strip the residue may include use ofadditional catalysts that are known to be effective for catalyticcracking, and the use of such stripper gas in combination with catalyticcracking catalysts is referred to herein as a strip-crack process. Thehydrogen is preferably either recycled in the reactor or compressed withthe acetol for a second reaction at higher pressure.

The preferred reaction process includes a first reactor. In the firstreactor, the first product and alternative product are removed as vaporeffluents from a liquid reaction where a sufficient hydrogen pressure ispresent to reduce the residue formation by at least 50% as compared tothe residue formation rate without hydrogen present. The preferredhydrogen partial pressures are between 0.2 and 50 bars, more preferablybetween 0.5 and 30 bars, and most preferably between 0.8 and 5 bars.

To achieve higher conversions to the alternative product, the firstproduct may be reacted in a second reactor that is operated at higherpartial pressures of hydrogen. In the second reactor, the partialpressure of hydrogen is at least twice the partial pressure of hydrogenin the second reactor, more preferably the partial pressure of hydrogenis at least four times the partial pressure of hydrogen in the firstreactor.

The respective temperatures of the first and second reactors arepreferably above the normal boiling point of the first product.

The use of hydrogen has an additional advantage of reducing residue thattends to deactivate catalysts which are useful in the disclosed process.In this sense, the hydrogen may be used as the gas purge or strippergas, as well as a reagent in the first reactor. For example, in thecracking of petroleum to gasoline, it is well-known that hydrogenreduces the formation of residue that tends to deactivate of thecatalyst; however, the use of hydrogen is more expensive than crackingof petroleum in the absence of hydrogen. In petroleum industry practice,considerable catalytic cracking is performed in the absence of hydrogenwith product loss, and specialized equipment is devoted to regeneratingthe deactivated catalyst. Those other practices differ from thepresently disclosed use of hydrogen stripper gas that is sufficient toreduce catalyst deactivation, but is sufficiently low in amount andamount/pressure to allow the non-hydrogen cat-cracking to dominate,e.g., at a pressure less that 50 bars, while the reaction is underway.

FIG. 14 shows one embodiment that implements these concepts. Processequipment 1400 the process where the hydrogen is compressed to proceedto the second reaction. In FIG. 14, like numbering is maintained withrespect to identical elements as shown in FIG. 13. The reaction processproceeds as described with respect to FIG. 13, except low pressurehydrogen stripper gas 1402 is applied to reactor separator 102, forexample, at a pressure slightly above atmospheric pressure. Althoughsome of this gas does result in the production of propylene glycol, thestripping of acetol is predominant. A mixture of acetol, propyleneglycol and water vapor flows through overhead line 402 to condenser1404, which condenses the vapors for use in the follow-on reactor 404.The extant hydrogen is optionally supplemented by additional hydrogen106 to establish the preferred reaction conditions discussed above.

FIG. 15 shows another embodiment, that of process equipment 1500. InFIG. 6, like numbering is maintained with respect to identical elementsas shown in FIG. 13. In process equipment 1500, the effluent throughintermediate overhead line 402 is applied to a series of condensers1502, 1504 that decrease in their relative temperatures to condensefirst the acetol in acetol condenser 1504 and then water in watercondenser 1504. The condensed acetol is applied to line 1506, e.g., bypumping at the requisite pressure, for delivery to the follow-on reactor404. Water effluent from water condenser 604 is discharged as waterpurge 1508.

The hydrogen pressures (or partial pressures) for this process in therespective reactor vessels may be lower than is required for “good”hydro-cracking and/or hydrogenolysis but sufficient to stop catalystdeactivation.

In addition to reactor configurations, other methods known in thescience for reducing residue (often an oligomer) formation is the use ofa solvent. The solvent can reduce residue formation or dissolve theresidue therein extending catalyst life. Solvents are preferably notreactive liquids. Supercritical solvents, such as carbon dioxide, havealso been demonstrated as effective for extending catalyst life whenresidue formation otherwise coats the catalyst.

Additional Examples of Reactor Performance

The following tables summarize additional process results from the pilotreactor.

TABLE 30 Effect of Temperature and Pressure on the Formation ofPropylene Glycol from Acetol. Pressure Log Reactor of [PG:Acetol Reactor[PG:Acetol Temperature Discharge Acetol PG Mass Temperature 1000/ Mass[° C.] [bar] [wt %] [wt %] Ratio] [K] T(K) Ratio] 202 1 20.14 36.94 1.83475.15 2.10 0.26 211 1 14.88 36.83 2.48 484.15 2.07 0.39 241 1 13.4321.1 1.57 514.15 1.94 0.20 241 1 12.74 20.37 1.60 514.15 1.94 0.20 201 113.17 53.15 4.04 474.15 2.11 0.61 240 1 30.18 29.81 0.99 513.15 1.95−0.01 242 1 31.76 33.19 1.05 515.15 1.94 0.02 239 1 24.54 26.38 1.07512.15 1.95 0.03 184 2 8.56 57.3 6.69 457.15 2.19 0.83 202 2 11.12 43.143.88 475.15 2.10 0.59 217 2 14.32 39.91 2.79 490.15 2.04 0.45 220 221.64 30.28 1.40 493.15 2.03 0.15 221 2 14.92 43.51 2.92 494.15 2.020.46 218 2 13.86 46.82 3.38 491.15 2.04 0.53 221 2 12.91 38.95 3.02494.15 2.02 0.48 243 2 44.84 41.83 0.93 516.15 1.94 −0.03 201 4 6.0355.15 9.15 474.15 2.11 0.96 203 4 4.68 60.46 12.92 476.15 2.10 1.11 2444 19.14 22.64 1.18 517.15 1.93 0.07 203 4 5.29 59.89 11.32 476.15 2.101.05 242 4 10.3 34.33 3.33 515.15 1.94 0.52 240 4 11.34 43.86 3.87513.15 1.95 0.59

TABLE 31 Equilibrium constant and Gibbs free energy values for thePropylene Glycol to Acetol Reaction. Reactor K eq, Temperature P_(TOTAL)YPG/ YPG/ YPG/ ΔG [K] [bar] YAC YPG YAC Y H₂ (YAC * YH2) (YAC * YH2 * P)[J/gmol] 450 2 0.06 0.85 13.55 0.89 17.84 7.58 −7579.35 454 2 0.11 0.807.52 0.89 10.58 4.21 −5430.3 457 2 0.11 0.82 7.73 0.89 10.59 4.33−5572.69 454 2 0.11 0.81 7.31 0.89 10.08 4.09 −5319.77 453 2 0.11 0.837.88 0.89 10.61 4.41 −5590.19 456 2 0.10 0.84 8.85 0.86 12.18 5.13−6197.74 480 2 0.15 0.72 4.69 0.86 7.63 2.74 −4020.2 476 1 0.18 0.643.61 0.74 7.68 4.91 −6296.26 475 1 0.21 0.68 3.30 0.74 6.52 4.44−5887.31 477 4 0.09 0.70 7.54 0.93 11.58 2.03 −2812.42 470 4 0.05 0.8315.78 0.92 20.60 4.27 −5675.43 493 2 0.18 0.60 3.39 0.85 6.58 1.99−2811.27 489 2 0.17 0.59 3.45 0.85 6.89 2.02 −2865.41 510 1 0.34 0.330.98 0.74 4.00 1.32 −1188.67 510 2 0.33 0.33 1.00 0.85 3.55 0.59 2243.48513 2 0.22 0.40 1.80 0.85 5.31 1.06 −250.565 515 2 0.21 0.34 1.63 0.855.61 0.96 158.9529 512 1 0.30 0.34 1.11 0.73 4.51 1.52 −1779.86 515 40.06 0.18 2.97 0.85 19.69 0.87 577.815 515 4 0.18 0.44 2.44 0.85 6.460.72 1427.368 514 4 0.08 0.36 4.43 0.92 13.55 1.21 −805.754 513 4 0.120.52 4.45 0.92 9.34 1.21 −814.523 513 4 0.13 0.51 3.97 0.91 8.47 1.09−362.552 512 4 0.22 0.57 2.59 0.92 4.94 0.71 1478.77

TABLE 32 Equilibrium constant and Gibbs free energy values for theAcetol to Propylene Glycol Reaction. Reactor K eq, Temperature P_(TOTAL)YPG/ YPG/ YPG/ ΔG [K] [bar] YAC YPG YAC Y H₂ (YAC * YH2) (YAC * YH2 * P)[J/gmol] 457 2 0.05 0.30 6.52 0.86 25.62 3.79 −5059.54 475 1 0.10 0.171.79 0.48 21.56 3.70 −5171.55 475 2 0.05 0.21 3.78 0.85 21.74 2.23−3172.27 474 1 0.07 0.28 3.93 0.75 18.54 5.25 −6534.36 476 4 0.03 0.3111.02 0.92 39.06 2.99 −4333.03 474 4 0.03 0.27 8.91 0.92 35.93 2.43−3496.03 476 4 0.02 0.31 12.58 0.92 44.10 3.43 −4879.46 484 1 0.07 0.182.41 0.72 19.24 3.37 −4888.07 490 2 0.06 0.17 2.71 0.87 18.49 1.56−1817.34 493 2 0.10 0.14 1.36 0.87 11.56 0.78 993.421 494 2 0.07 0.202.84 0.87 16.25 1.64 −2022.76 491 2 0.06 0.21 3.29 0.87 18.02 1.89−2607.56 494 2 0.06 0.18 2.94 0.84 19.18 1.74 −2272.33 516 2 0.18 0.160.91 0.84 6.72 0.54 2634.137 513 1 0.14 0.13 0.96 0.64 11.44 1.50−1725.2 515 1 0.13 0.14 1.02 0.71 10.53 1.43 −1536.74 514 1 0.06 0.091.53 0.71 22.71 2.15 −3267.45 514 1 0.06 0.10 1.56 0.71 22.91 2.19−3342.4 512 1 0.12 0.12 1.05 0.72 11.85 1.45 −1576.73 515 4 0.05 0.163.25 0.91 22.41 0.90 472.8089 513 4 0.06 0.21 3.77 0.92 19.43 1.03−110.795 517 4 0.12 0.14 1.15 0.92 9.05 0.31 4975.326

TABLE 33 Effect of temperature and pressure on total byproducts of theglycerol to propylene glycol reaction. Pressure of Reactor Total ReactorDischarge Temperature Byproducts Temperature [° C.] [bar] [K] [wt %] 2201 493 6.29 220 1 493 15.50 238 1 511 45.29 241 1 514 42.06 240 1 51321.31 220 2 493 4.27 220 2 493 4.98 220 2 493 5.37 221 2 494 11.36 220 2493 18.65 221 2 494 16.46 237 2 510 22.32 236 2 509 22.01 240 2 51329.03 220 4 493 4.54 220 4 493 1.58 240 4 513 17.59

TABLE 34 Effect of the residence time on byproducts of the glycerol topropylene glycol reaction. Reactor Pressure of Hydrogen Reactor TotalWater Temperature Discharge flowrate Volume Residence time ByproductsContent [° C.] [bar] [(ft{circumflex over ( )}3)/min] [ft{circumflexover ( )}3] [min] [wt %] [wt %] 220 1 0.59 0.05 0.08 6.29 18.52 220 10.78 0.05 0.06 15.5 21.96 220 2 1.28 0.05 0.04 4.27 18.93 220 2 0.870.05 0.06 4.98 13.71 220 2 0.87 0.05 0.06 5.37 13.30 221 2 0.18 0.050.28 11.36 20.74 220 2 0.18 0.05 0.28 18.65 21.01 221 2 0.18 0.05 0.2816.46 23.29 220 4 1.06 0.05 0.05 4.54 17.16 220 4 0.78 0.05 0.06 1.5815.73 238 1 0.09 0.05 0.56 45.29 15.72 241 1 0.09 0.05 0.56 42.06 14.02240 1 0.18 0.05 0.28 21.31 25.70 237 2 0.18 0.05 0.28 22.32 24.01 236 20.18 0.05 0.28 22.01 24.09 240 2 0.18 0.05 0.28 29.03 26.63 240 4 0.180.05 0.28 42.86 28.78 240 4 0.35 0.05 0.14 17.59 26.51

TABLE 35 Effect of the water content on byproducts of the glycerol topropylene glycol reaction Pressure of Water Total Reactor Dischargecontent Byproducts Temperature [° C.] [bar] [wt %] [wt %] 220 1 18.526.29 220 1 21.96 15.5 240 1 25.7 21.31 220 2 18.93 4.27 220 2 13.71 4.98220 2 13.3 5.37 221 2 20.74 11.36 220 2 21.01 18.65 221 2 23.29 16.46237 2 24.01 22.32 236 2 24.09 22.01 240 2 26.63 29.03 220 4 17.16 4.54220 4 15.73 1.58 240 4 26.51 17.59

TABLE 36 Effect of temperature and pressure on byproducts of thepropylene glycol to acetol reaction. Reactor Pressure of TotalTemperature Discharge Reactor Byproducts [° C.] [bar] Temperature [K][wt %] 203 1 476 18.94 202 1 475 10.41 237 1 510 27.84 239 1 512 21.51177 2 450 3.13 178 2 451 5.11 184 2 457 4.02 181 2 454 4.53 182 2 4553.99 183 2 456 3.35 207 2 480 10.84 220 2 493 16.62 216 2 489 18.99 2372 510 25.22 240 2 513 27.74 242 2 515 35.03 204 4 477 8.79 197 4 4701.82 242 4 515 30.03 242 4 515 19.7 241 4 514 39.39 240 4 513 21.98 2404 513 25.35 239 4 512 16.43

TABLE 37 Effect of the residence time on byproducts of the propyleneglycol to acetol reaction. Pressure Reactor of Hydrogen Hydrogen ReactorTotal Temperature Discharge flowrate flowrate Volume ResidenceByproducts [° C.] [bar] [L/min] [(ft{circumflex over ( )}3)/min][ft{circumflex over ( )}3] time [min] [wt %] 177 2 5.0 0.2 0.05 0.283.13 178 2 5.0 0.2 0.05 0.28 5.11 184 2 5.0 0.2 0.05 0.28 4.02 181 2 5.00.2 0.05 0.28 4.53 182 2 5.0 0.2 0.05 0.28 3.99 183 2 5.0 0.2 0.05 0.283.35 203 1 2.5 0.1 0.05 0.56 18.94 202 1 5.0 0.2 0.05 0.28 10.41 207 25.0 0.2 0.05 0.28 10.84 204 4 2.5 0.1 0.05 0.56 8.79 197 4 5.0 0.2 0.050.28 1.82 220 2 5.0 0.2 0.05 0.28 16.62 216 2 5.0 0.2 0.05 0.28 18.99239 1 2.5 0.1 0.05 0.56 21.51 237 1 5.0 0.2 0.05 0.28 27.84 237 2 5.00.2 0.05 0.28 25.22 240 2 5.0 0.2 0.05 0.28 27.74 242 2 5.0 0.2 0.050.28 35.03 242 4 5.0 0.2 0.05 0.28 30.03 242 4 5.0 0.2 0.05 0.28 19.7241 4 5.0 0.2 0.05 0.28 39.39 240 4 10.0 0.4 0.05 0.14 21.98 240 4 10.00.4 0.05 0.14 25.35 239 4 2.5 0.1 0.05 0.56 16.43

TABLE 38 Effect temperature and pressure on byproducts of the acetol topropylene glycol reaction. Reactor Reactor Temperature Pressure ofTemperature Total Byproducts [° C.] Discharge [bar] [K] [wt %] 202 1 4758.12 211 1 484 12.63 241 1 514 34.39 241 1 514 38.74 201 1 474 4.48 2401 513 17.74 242 1 515 29.39 239 1 512 16.08 184 2 457 2.94 202 2 4756.11 217 2 490 13.58 220 2 493 14.83 221 2 494 12.56 218 2 491 12.27 2212 494 14.79 243 2 516 29.96 201 4 474 5.2 203 4 476 4.04 244 4 517 41.81203 4 476 6.63 242 4 515 11.84 240 4 513 12.77

TABLE 39 Effect of the residence time on byproducts of the acetol topropylene glycol reaction. Pressure Reactor of Hydrogen Hydrogen ReactorTotal Temperature Discharge flowrate flowrate Volume ResidenceByproducts [° C.] [bar] [L/min] [(ft{circumflex over ( )}3)/min][ft{circumflex over ( )}3] time [min] [wt %] 184 2 5.0 0.18 0.05 0.282.94 202 1 2.5 0.09 0.05 0.56 8.12 201 1 5.0 0.18 0.05 0.28 4.48 202 25.0 0.18 0.05 0.28 6.11 201 4 2.5 0.09 0.05 0.56 5.20 203 4 2.5 0.090.05 0.56 4.04 203 4 5.0 0.18 0.05 0.28 6.63 211 1 2.5 0.09 0.05 0.5612.63 217 2 5.0 0.18 0.05 0.28 13.58 220 2 5.0 0.18 0.05 0.28 14.83 2212 5.0 0.18 0.05 0.28 12.56 218 2 5.0 0.18 0.05 0.28 12.27 221 2 5.0 0.180.05 0.28 14.79 241 1 2.5 0.09 0.05 0.56 34.39 241 1 2.5 0.09 0.05 0.5638.74 240 1 5.0 0.18 0.05 0.28 17.74 242 1 5.0 0.18 0.05 0.28 29.39 2391 10.0 0.35 0.05 0.14 16.08 243 2 5.0 0.18 0.05 0.28 29.96 244 4 2.50.09 0.05 0.56 41.81 242 4 5.0 0.18 0.05 0.28 11.84 240 4 10.0 0.35 0.050.14 12.77

Additional Examples of Effective Hydrogenation Catalysts

Most of the experiments validating the embodiments of these inventionswere demonstrated using a copper chromite catalyst. Table 40 summarizesa survey of additional catalysts for the hydrogenation of acetol topropylene glycol.

TABLE 40 Effect of the residence time on byproducts of the acetol topropylene glycol reaction. The reactions were at 185° C., 14 bars ofhydrogen overpressure, and the reaction time was 4 hours. Acetolconversion Propylene Glycol Supplier Description (%) Selectivity (%)Grace Davison Raney Copper 99.07 91.72 Degussa 5% Palladium/Carbon 76.2274.26 Sud-Chemie Copper-Zinc a 91.56 87.17 Sud-Chemie Copper/Alumina b82.67 96.91 Sud-Chemie Copper Chromium c 96.89 98.92 promoted by Ba andMn Sud-Chemie Copper Chromium d 98.22 93.86 Sud-Chemie Copper Chromium74.22 95.97 promoted by Ba e Engelhard Copper Chromium 98.00 96.08promoted by Mn f In-house Copper/Silica g 82.67 93.67 Grace DavisonRaney Nickel 99.56 98.90 Degussa 5% Platinum/Carbon 72.89 88.71 JohnsonMatthey 5% Ruthenium/Carbon 100.00 100.00 Alfa AesarNickel/silica-alumina 73.78 81.20 Johnson Matthey Nickel/Carbon 90.2289.16 Nominal Compositions (wt %): a CuO (33), ZnO (65), Al2O3 (2) b CuO(56), Al2O3 (34), MnO2 (10) c CuO (45), Cr2O3 (47), MnO2 (3.5), BaO(2.7) d CuO (50), Cr2O3 (38) e CuO (41), Cr2O3 (46), BaO (13) f CuO(36), Cr2O3 (33), MnO2 (3) g 23 wt % copper on silica supportApplicability to Broader Reaction Mechanisms

The process that has been shown and described has been proven effectivefor production of acetol and propylene glycol, but is not limited to thereaction mechanisms of FIGS. 2 and 3. The process and process equipmentis generally applicable to a range of reactions having similar overallmechanisms including at least four classes of such reactions in contextof the discussion below.

A first class of liquid phase catalytic reaction occurs where a reactant(e.g. glycerol) distributes predominantly in a liquid phase and thereactant is converted to at least a first product (e.g. acetol) thatthat has a boiling point at least 20° C. lower in temperature than thereactant.

A second class of liquid phase catalytic reaction occurs where thereactant reacts in a parallel mechanism with hydrogen to form at leastone alternative product (e.g. propylene glycol) where the alternativeproduct has a boiling point that is at least 20° C. lower in temperaturethan the reactant. The selectivity to formation of the alternativeproduce(s) from this second reaction is greater than 0.5 when in thepresence of hydrogen and hydrogen partial pressures in excess of 100bars.

A third class of reaction proceeds substantially in parallel the firstreaction including the reactant forming a higher molecular weightresidue species that directly or indirectly reduces the effectiveness ofthe catalyst promoting the first reaction.

A fourth class of reaction that occurs when hydrogen is present thatsubstantially inhibits the formation of the residue of the thirdreaction where the rate of formation of residue is reduced by at least50% with the hydrogen partial pressure in 50 bars.

The process includes use in appropriate reactor configurations, such asthe process equipment discussed above.

Those skilled in the art will appreciate that the foregoing discussionteaches by way of example, not by limitation. The disclosedinstrumentalities set forth preferred methods and materials, and may notbe narrowly construed to impose undue limitations on the invention. Thescope of the inventor's patentable inventions is defined by the claims,nothing else. Furthermore, the inventors hereby state their intention torely upon the Doctrine of Equivalents to protect the full scope of theirrights in what is claimed.

1. A process for converting glycerol to a product at high selectivity topropylene glycol and low selectivity to ethylene glycol, comprising:providing a gas phase reaction mixture that is essentially free ofliquid and contains glycerol with a partial pressure of glycerol in arange from 0.01 bars and 0.5 bars of glycerol, and hydrogen, thereaction mixture being maintained at a total pressure between 0.02 and10 bars; and contacting the gas phase reaction mixture with aheterogeneous catalyst at a temperature between 150° C. and 280° C. toform propylene glycol.
 2. The process of claim 1, wherein theheterogeneous catalyst used in the step of contacting contains at leastone element from Groups I, VI, and VIII of the Periodic Table.
 3. Theprocess of claim 1, wherein the heterogeneous catalyst used in the stepof reacting contains at least one metal selected from the groupconsisting of copper and chromium.
 4. The process of claim 3, whereinthe heterogeneous catalyst is a copper chromite catalyst.
 5. The processof claim 1, wherein the temperature of the contacting step is within arange from 180° C. to 240° C.
 6. The process of claim 1, wherein thepressure of the contacting step is within a range from 1 bar to 10 bars.7. The process of claim 1, wherein a reaction product formed from thereaction mixture contains acetol.
 8. The process of claim 7, whereinacetol is removed from the reaction product and recycled to form amixture with the glycerol feed.
 9. The process of claim 1 including astep of contacting a glycerol containing feed with a gas in anevaporating step at a temperature between 150° C. and 280° C. toevaporate glycerol to form a reactor feed.
 10. The process of claim 9wherein the contacting of glycerol with the gas occurs at a temperaturewithin a range from 200° C. to 260° C.
 11. The process of claim 10wherein the temperature range is from 210° C. to 230° C.
 12. The processof claim 9 wherein the glycerol feed contains a non-volatile salt andthe evaporating step partially separates the non-volatile salt from theglycerol.
 13. The process of claim 9 including the step of recycling thegas through a condenser that separates the reaction product from thegas.
 14. The process of claim 1 wherein reaction conditions of thecontacting step include a partial pressure of glycerol being less thanglycerol's dew point partial pressure in the reaction mixture, andgreater than one fourth the dew point partial pressure in the reactionmixture.
 15. The process of claim 14 wherein the partial pressure ofglycerol is greater than half the dew point partial pressure in thereaction mixture.
 16. The process of claim 1 wherein the reactionconditions of the contacting step includes the temperature being lessthan 230° C.
 17. The process of claim 1 further comprised of at leasttwo reaction steps where a first reaction step has reaction temperaturesgreater than 200° C. and the second reaction step has temperatures lessthan 220° C. and where the glycerol concentration entering the secondreaction step is less than half the glycerol concentration entering thefirst reaction step.
 18. The process of claim 17 wherein the firstreaction step has reaction temperatures greater than 210° C. and thesecond reaction step has temperatures less than 210°0 C.
 19. A processfor converting glycerol to a product at high selectivity to propyleneglycol and low selectivity to ethylene glycol, comprising: contacting agas phase reaction mixture with a heterogeneous catalyst, wherein thegas phase that contains essentially no liquid and has a partial pressureof glycerol between 0.01 and 0.5 bars of glycerol and a total pressurebetween 1 and 15 bars; and establishing a temperature between 150° C.and 280° C. to facilitate a reaction.
 20. The process of claim 19,wherein the heterogeneous catalyst used in the step of contactingcontains at least one element from Groups I or VIII of the PeriodicTable.
 21. The process of claim 18 where the partial pressure ofglycerol is less than glycerol's dew point partial pressure in thereaction mixture and greater than one fourth the dew point partialpressure in the reaction mixture.
 22. The process of claim 21 where thepartial pressure of glycerol is greater than half the dew point partialpressure in the reaction mixture.
 23. A process for converting glycerolto a product at high selectivity to a mixture of acetol and propyleneglycol and low selectivity to ethylene glycol, comprising: contacting agas phase reaction mixture with a heterogeneous catalyst, wherein thegas phase reaction mixture contains essentially no liquid and has apartial pressure of glycerol between 0.01 and 0.5 bars of glycerol and apartial pressure of hydrogen between 0.01 and 5 bars of hydrogen and atotal pressure between 0.02 and 5 bars; and establishing a temperaturebetween 150° C. and 250° C. to facilitate a reaction.
 24. The process ofclaim 23, wherein the heterogeneous catalyst used in the step ofcontacting contains at least one element from Groups I or VIII of thePeriodic Table.
 25. The process of claim 23, wherein the total pressureof reaction between 0.02 and 25 bars.
 26. The process of claim 25 wherethe partial pressure of glycerol is less than glycerol's dew pointpartial pressure in the reaction mixture and greater than one fourth thedew point partial pressure in the reaction mixture.
 27. The process ofclaim 26 where the partial pressure of glycerol is greater than half thedew point partial pressure in the reaction mixture.
 28. The process ofclaim 1 wherein the reaction mixture is maintained at a total pressurebetween 0.02 and 5 bars.
 29. The process of claim 28 wherein the partialpressure of hydrogen is maintained at a total pressure between 0.01 and5 bars.
 30. The process of claim 1 wherein the reaction mixture ismaintained at a total pressure between 1 and 5 bars.
 31. The process ofclaim 1 wherein the reaction mixture is maintained at a total pressurebetween 0.5 and 3 bars.
 32. The process of claim 19 wherein the reactionmixture is maintained at a total pressure between 1 and 5 bars.
 33. Theprocess of claim 23 wherein the reaction mixture is maintained at atotal pressure between 1 and 5 bars.
 34. The process of claim 23 whereinthe reaction mixture is maintained at a total pressure between 0.5 and 3bars.